DE69328869T2 - GAS SEPARATION AT TEMPERATURES UNDER ROOM TEMPERATURE USING MEMBRANES MADE OF GLASS-LIKE POLYMERS - Google Patents

Description

BACKGROUND OF THE INVENTION

The invention relates to a method for separating a gas mixture, which uses a membrane with a separation layer or a separation region of a glassy polymer, which is operated at temperatures below the ambient temperature. The present invention further relates to a device for such a method.

The use of membranes to separate different components of gas mixtures is well known. Membranes have been used to separate, remove, purify or partially recover a variety of gases, including hydrogen, helium, oxygen, nitrogen, argon, carbon monoxide, carbon dioxide, ammonia, water vapor, methane and other light hydrocarbons.

Membrane separations are based on the relative permeability of two or more component gases through the membrane. In order to separate a gas mixture into two sections, a section richer and poorer in at least one component gas, the gas mixture is brought into contact with a side of the membrane through which at least one of the component gases selectively permeates. A component gas that selectively permeates through the membrane penetrates the membrane more quickly than at least one other component gas of the gas mixture. The gas mixture is thereby divided into a stream that is enriched in the selectively permeating component gas(es) and a stream that is depleted in the selectively permeating component gas(es). A relatively non-permeating component gas penetrates the membrane to a lesser extent than at least one other component gas of the gas mixture. A suitable membrane material is chosen such that a certain degree of separation of the gas mixture can be achieved.

D. W. Brubaker et al., Industrial and Engineering Chemistry, Volume 46, pages 1465-1473 (1953), discloses gas permeation through various films and parameters affecting such permeation. R. W. Roberts, Journal of Applied Polymer Science, Volume 7, No. 6, pages 2183-2197 (1963), investigated the effect of temperature on permeation through a variety of polymer films.

S. Srinivasan presented a paper entitled "An Extraordinary Polymeric Membrane That Rejects Light Gases" at the Gordon Research Conference on Synthetic Membranes on July 10, 1990, in which he presented permeability data on a polytrimethylsilylpropyne membrane at temperatures down to -40ºC for nitrogen and carbon dioxide. Anomalous gas mixture effects were observed. These effects were explained by proposing that transport in this material was dominated by diffusion in pores and on the surfaces of these pores, rather than the conventional solution-diffusion mechanism. Thorogood presented permeability data for oxygen and nitrogen at low temperatures using a glassy film of density-modified polytrimethylsilylpropyne at the International Gas Separation Meeting in Austin, Texas on April 23, 1991. See also S. R. Auvil et al., Book of Abstracts, The Fourth Chemical Congress of North America, 25-30. August 1991, point 119.

D. J. Moll et al., at "The First Annual National Meeting of the North American Membrane Society" on June 3-5, 1987, disclosed permeability data for helium, argon, and xenon using tetramethyl bisphenol A polycarbonate films over a wide temperature range.

D. J. Moll et al., at "The Gordon Research Conference on Reverse Osmosis, Ultrafiltration, and Gas Separation" on July 31, 1989, disclosed permeability data for hydrogen and deuterium using a film of bisphenol A polycarbonate at temperatures down to -125ºC.

In general, the gas permeability of known membrane materials decreases with decreasing temperature. In general, the prior art recommended that membrane separations be carried out at temperatures that maintain good physical properties of the membrane and avoid condensation of the gaseous components to be separated. Typically, the temperature is kept as high as possible without deleterious effects on the physical integrity or performance of the membrane, since higher temperatures generally increase the rate of gas permeation through the membrane. Temperatures in the range of 20ºC to 40ºC are generally used in prior art separations. For separations of many gas mixtures using membranes, conditions near ambient temperature are suitably used.

In separations of specific processes from certain gas mixtures, temperatures of less than zero degrees Celsius have occasionally been used. US Patent 4,793,829 describes the selective permeation of ammonia through a polysulfonamide membrane at a temperature in the range of zero degrees Celsius to -20ºC. Canadian Patent 1,188,630 describes the separation of carbon dioxide from methane and other light hydrocarbons using a cellulose ester membrane at temperatures of less than 10ºC, preferably between -15ºC and 5ºC. See also D. Parro, Energy Process, Volume 6, No. 1, pages 51-54 (1985), D. Parro, Technology, Oil & Gas Journal, September 24, 1984, pages 85-88, and G. Cutler et al. al., Laurance Reid Gas Conditioning Conference, 4-6. March 1985.

Anand et al., U.S. Patent 4,840,646, describes the permeation, selectivity and flux numbers obtained at 24°C for polyester copolymers. These polymers disclosed by Anand are excluded from the present invention. Figure 7 here graphically shows the polymers and copolymers that meet claim 1. This graph was prepared using permeation selectivity and flux data obtained at 30°C.

Research is continuing into the development of new membranes that have high separation factors and high permeation rates for selected component gases of gas mixtures of interest. However, the use of temperatures below 20ºC in membrane separations is generally viewed with disapproval. Temperatures of zero degrees Celsius are generally considered possible only for a limited number of specific gas mixtures and special membranes.

In general, there is a trade-off between separation factor and permeability when membranes are used to separate components of gas mixtures. Membranes with a high separation factor generally have a lower permeability than membranes with a lower separation factor. This trade-off is very clear when separating mixtures of oxygen and nitrogen and mixtures of carbon dioxide and methane.

Typically, the permeability of gases through membranes decreases rapidly with temperature. The separation factor generally increases with decreasing temperature, but the rate of change of the separation factor with temperature depends on the specific chemical properties of the membrane and the component gases of the gas mixture to be separated.

In the case of oxygen/nitrogen separations, both the solubility in and the diffusivity of the gases through the membrane affect the separation performance. In general, oxygen dissolves in most polymers to a concentration 50 to 100 percent higher than the concentration of nitrogen. While the solubility component of selectivity is important in the separation of oxygen and nitrogen, the diffusivity component of selectivity is the primary factor in determining overall selectivity.

The diffusivity depends on the dimension, which is related to the repulsive component of the interaction between gas molecules and the material forming the membrane. In simplified form, this depends on the size of the molecules diffusing through the material, such as a polymer. Smaller gas molecules can enter and diffuse through a polymer matrix more easily. In general, molecules diffuse more easily through a polymer with a larger unoccupied volume, i.e., "free volume". Accordingly, in a polymer with relatively large free volumes, the diffusivity of gases is generally higher than in a polymer with smaller free volume. However, this more open structure increases the diffusivity of both gases and leads to lower selectivity.

Although unoccupied volume in a polymer is important for determining membrane separation properties, other factors are very important for achieving improvements in such properties. The dynamics of the movement of the polymer chains, the local movement of the repeating units and the vibrations of smaller subunits can determine the have a detrimental effect on separation properties. These movements prevent the polymer from maintaining an optimal spacing between subunits of the polymer for the desired separation. In general, lower temperatures reduce the frequency and amplitude of movements of the polymer matrix.

While lower temperatures may limit the movement of the polymer matrix, temperatures much below 20ºC have not generally been used for membrane separation of gases. The reluctance of the art to consider subambient temperatures for membrane separation of gases arises due to problems related to phase changes in gas mixture components, the higher operating costs resulting from cooling feed gas mixtures, and the reduced permeation rate at lower temperatures.

A method of separating component gases of a gas mixture which achieves a higher than normal separation factor has been achieved and acceptable permeability is still sought. A method of operation which improves the performance of existing membrane materials and potentially increases the performance of new membrane materials is desirable.

SUMMARY OF THE INVENTION

The invention relates to a method of separating a permeate gas stream and a non-permeate gas stream from a gas mixture. The gas mixture contacts a first side of a gas separation membrane. This gas separation membrane has a polymeric separation layer or a polymeric separation region which, at the separation temperature, a glassy state. A chemical potential difference is maintained between the first side of the membrane and the second side of the membrane such that at least a first component gas present in the gas mixture selectively permeates through the membrane from the first side to the second side of the membrane relative to at least a second component gas present in the gas mixture such that the permeating gas stream is enriched in the first component gas. Contact between the gas mixture and the membrane occurs at a temperature at or above the freezing point of any liquid present on or within the membrane. Thus, when pure water is present, the operating temperature is selected from the range between 1ºC and 5ºC. When salts are present, the freezing point is depressed and the permeation temperature is between 1ºC and 5ºC above this depressed freezing point. The membrane is chosen such that when a mixture of 80 mole percent nitrogen and 20 mole percent oxygen is used as feed at 30°C with a pressure of 30 psia on the first side of the membrane and a vacuum of less than 1 mm Hg on the second side of the membrane, the permeability of oxygen in Barrers is less than 2000 and has the following relationship to the selectivity for oxygen relative to nitrogen (hereinafter referred to as Equation I):

Permeability > 2000/(selectivity)7/2

Selectivity in the sense of formula I is defined as the ratio of the permeability of oxygen to the permeability of nitrogen. The first side and the second side of the membrane are separated and not connected, so that it is possible to separate the gas that does not permeate through the membrane, that is, the rejected or non-permeating Gas, and/or the gas which permeates through the membrane, i.e. the permeating gas.

According to the invention, a method for separating component gases in a gas mixture is provided, comprising:

A/ Contacting a first side of a gas separation membrane, the membrane comprising a polymeric separation layer or region which is in a glassy state under the conditions in which the separation is carried out, with a gas mixture while maintaining a difference in chemical potential from the first side of the membrane to a second side of the membrane such that at least a first component gas of the gas mixture selectively permeates to at least a second component gas in the gas mixture from the first side of the membrane to the second side of the membrane, said contacting occurring at a temperature of 5°C or below, provided that the permeation occurs at or above the freezing point of any liquid present, the membrane being selected such that when using a mixture of 80 mole percent nitrogen and 20 mole percent oxygen as the feed at 30°C at a pressure of 205.5 kPa (30 psia) on the first side of the membrane and a vacuum of less than 0.133 kPa (1 mm Hg) on the second side of the membrane, the permeability of oxygen in Barrers is less than 2000 and has the following relationship for oxygen/nitrogen selectivity:

Permeability > 2000/(selectivity)7/2; and

B/ recovering at least one of the permeate gas or the rejected non-permeate gas, wherein it/this Gas separation-producing layer/region is a glassy polymeric composition of material having a repeating unit of the formula:

where U is independently selected from chlorine or bromine.

A certain novel apparatus for carrying out the method of the invention is also an embodiment of the invention.

Such a device using the previously defined membrane forms the subject matter of claims 41-56.

It is believed that the process of the present invention achieves significantly improved performance compared to membrane separation processes at ambient temperature. With respect to air separation for the production of nitrogen, this improved performance can be characterized by a significantly higher recovery of nitrogen, that is, the production of nitrogen as a fraction of the feed air, at the same product purity, or the production of significantly higher purity product at the same recovery. Alternatively, it may be possible to produce nitrogen with similar purity and recovery and significantly higher productivity. In other separations, such as the separation of carbon dioxide and a light hydrocarbon such as methane, improved performance may also be characterized by higher recovery, higher productivity and/or higher purity of desired products. Improved performance may further be characterized by similar recovery and purity of product with smaller membrane surface area and/or lower power consumption.

Currently, cryogenic air separation is considered to be the most economical for large scale nitrogen production, ambient membrane separation is the most economical for small scale production, and pressure swing adsorption (PSA) is the most economical for medium scale production. These different areas of economic superiority arise from the varying importance of capital and operating costs for each conventional process at different levels of production; operating costs are more important for large scale production than for small scale production, and capital expenditure is more important for small scale production than for small scale production. Cryogenic air separation produces nitrogen at the lowest operating cost, but capital costs are high for small scale production. Ambient membrane separation produces nitrogen at the lowest capital cost, but operating costs are high for large scale production. The subambient temperature membrane process of the present invention is believed to have lower operating costs in compared to ambient air membrane separation. Capital costs could be higher due to the additional equipment required for cooling. However, under certain circumstances it is possible that the improved efficiency, i.e. selectivity and/or productivity, of the membrane system may enable a significant reduction in capital costs with respect to the compressor or membrane, respectively. This reduction may offset the additional capital required for cooling. In general, the process of the invention is expected to be more economical at a level of production which is higher than the level at which ambient temperature membrane separations are currently used. With respect to air separation, the process of the present invention is expected to be competitive with respect to larger scale PSA separations and smaller scale cryogenic separations. It is also expected that the invention is competitive with respect to other types of separations, for example separations involving CO₂, H₂S, SOx or other gases. More specifically, it is believed that the present invention can compete with separation technology based on amine contactor processes that use monoethanolamine, diethanolamine or other formulated amines. In some situations, it may be advantageous to combine the present invention with adsorption processes based on amines or other chemicals to form a hybrid system.

SHORT DESCRIPTION OF THE DRAWING

Fig. 1 shows in schematic form an apparatus suitable for achieving the process of the subject invention which applies an expansion device to the non-permeate stream.

Fig. 2 shows in schematic form an apparatus suitable for achieving the process of the subject invention which applies an expansion device to the permeate stream.

Fig. 3 shows in schematic form an apparatus suitable for achieving the method of the subject invention which uses an external freon cooling system for cooling the feed gas mixture.

Fig. 4 shows in schematic form an apparatus suitable for achieving the process of the subject invention, which uses an external freon cooling system for cooling and membranes for drying the feed gas mixture.

Fig. 5 shows in schematic form a device suitable for achieving the method subject of the invention, which uses a compressor operated by an expansion device.

Fig. 6 shows in schematic form a device suitable for achieving the process object of the invention, which uses a vacuum pump on the permeate stream.

Fig. 7 shows in schematic form a device suitable for achieving the process object of the invention, which uses a cascaded device and a recycle stream.

Fig. 8 shows in schematic form a device which is suitable for an alternative embodiment of the process of the invention which applies an expansion device to the non-permeate stream.

Figure 9 shows in schematic form an apparatus suitable for achieving an alternative embodiment of the method of the invention which uses an external Freon cooling system for cooling the feed gas stream.

Fig. 10 shows in schematic form an apparatus for achieving an alternative embodiment of the invention which uses a Joule-Thomson expansion of gas permeating through the membrane for cooling.

Fig. 11 shows an illustrative embodiment of a device in schematic form which is suitable for achieving the method which is the subject of the invention.

Figure 12 shows a comparison at different temperatures of the oxygen/nitrogen separation factor and oxygen permeability for two specific membranes using air as the feed gas. Membrane 1 illustrates the performance of a membrane suitable for the practice of the process of the present invention. Membrane 2 is a comparison of the performance of a membrane not illustrative of the present invention.

Figure 13 shows certain polymeric membrane materials that meet the separation properties of Equation I (materials labeled 1 through 6 and 10) and other polymeric membrane materials that did not meet these criteria (materials labeled 7 through 9).

Fig. 14 shows in schematic form an alternative apparatus suitable for achieving the process which is the subject of the invention when a high pressure permeate is desired.

Fig. 15 shows in schematic form an alternative device for producing a high-pressure permeate of higher purity.

Fig. 16 shows in schematic form an alternative device for generating a high-pressure non-permeate stream, in which cooling is carried out by an external cooling device.

For simplicity, no operations for pretreating the feed gas mixture, such as drying or filtering, are shown in Figures 8, 9, 10, 14, 15 or 16. In some embodiments, the feed gas mixture is advantageously dried prior to contacting the membrane so that water is not present in a significant amount, which may otherwise condense and freeze at the operating temperature of the membrane separation. In some embodiments, other condensable fluids are also advantageously removed from the feed gas mixture prior to contacting the membrane. Particles, oil and other contaminants that could adversely affect the physical integrity or performance of the membrane or other system components are also advantageously removed by conventional means.

DETAILED DESCRIPTION OF THE INVENTION

As defined in this description:

The "temperature" of permeation is critical and can be at or slightly above the freezing point of any liquid present on or within the membrane. The key feature is that liquid, such as water, does not freeze or crystallize on or within the membrane and thus does not affect or destroy the permeation characteristics thereof. That is, if salt water (e.g. sea water) is present, the freezing point of the liquid is lowered to below 0ºC, for example -1, -2, -3, -4, -5ºC or more below the freezing point. Preferably, permeation takes place at 2ºC higher than the freezing point of the liquid present, more preferably 1ºC higher than the freezing point of any liquid present.

It has now been found that certain membranes comprising separation layers or separation regions and made from glassy polymers exhibit dramatic improvements in the separation factors for specific gas pairs at decreasing temperatures. At the same time, there is a relatively small loss of gas permeability with decreasing temperature. In order to achieve the advantages of this process, both the separation layer and the separation region of the membrane as well as the component gases to be separated must meet certain criteria, which are set out below.

It has been found that by selecting a suitable membrane and gas mixture, membrane separations at reduced temperature have improved separation performance. In general, preferred membranes and gas pairs operate at temperatures of preferably -5ºC or less, more preferably about -25ºC or less, resulting in performance that is generally higher than membranes of known polymers operated at ambient temperatures. This improved membrane performance is typically demonstrated by higher selectivity at comparable or slightly reduced gas permeability for the more permeable gas of a gas mixture compared to the membrane at ambient temperature, i.e., 20ºC to 40ºC. This may also allow the advantageous use of membranes previously classified as having lower separation properties at ambient temperature.

The membrane includes a separation layer or region which is in a glassy state under the conditions under which the separation is conducted. Glassy polymers used herein belong to a class of materials well known in the art which are either non-crystalline or partially crystalline polymers which have a glass transition temperature above the temperature at which the membrane is used. Polymers used in this invention must have a glass transition temperature (Tg) which is at least 1°C, preferably at least 10°C above the operating temperature at which the membrane is used. Polymers useful in the present invention more preferably have a glass transition temperature of at least about 100°C, even more preferably at least about 150°C.

Particularly preferred membranes for use in the present invention have an increased oxygen/nitrogen separation factor at a temperature of 5°C or less for oxygen, i.e., the more permeable gas, relative to nitrogen of at least about 25 percent, more preferably at least about 50 percent, relative to the separation factor at 30°C. At the temperature at which the oxygen/nitrogen separation factor is increased by 25 percent relative to equal to the separation factor for oxygen/nitrogen at 30ºC, the permeability of oxygen decreased preferably by no more than about 50 percent, more preferably no more than about 40 percent, relative to the permeability at 30ºC.

One criterion for selecting a glassy polymer suitable for the separation layer or region of a membrane is to determine the change in the separation factor or selectivity, α, with respect to temperature, i.e., (dα/dT), and to compare this with the change in the gas permeability of the more permeable component gas, P, with respect to the separation factor. A useful formula for comparing membrane performance (hereinafter referred to as Equation II) is as follows:

Permselectivity index = ln(P) * (dα/dT)dln(P)/dα

The values and derivatives of gas permeability and separation factor in Equation II are calculated at 30°C. With respect to oxygen/nitrogen separation, the membranes useful in the present invention have a permselectivity index of preferably at least about 0.25, more preferably at least about 0.35, even more preferably at least about 0.40. Membranes having the properties required by Equation I generally have a desirable permselectivity index.

Further sets of preferred membranes for use in the present invention can be defined by the figure of merit. The figure of merit, defined for a particular gas pair, takes into account the effect of a variety of factors, such as the permeability of the more permeable gas, the change The effect of the rate of change of permeability with respect to the rate of change of selectivity, and the rate of change of selectivity with respect to the rate of change of temperature, on the performance of the membrane. The figure of merit for an oxygen/nitrogen separation is calculated by equation III.

where PO2 is the permeability of oxygen and αO2/N2 is the selectivity of oxygen/nitrogen. The values and derivatives are calculated at 30°C. The glassy polymers useful in the present invention have a figure of merit for oxygen/nitrogen separation for membranes made therefrom of at least about 1.0, more preferably at least about 3.0.

The figure of merit for a carbon dioxide/methane separation is calculated by equation IV:

where PCO₂ is the permeability of carbon dioxide and αCO₂/CH₄ is the permeability of carbon dioxide/methane. The values and derivatives are calculated at 30°C. The glassy polymers useful in the present invention have a figure of merit for carbon dioxide/methane separation for membranes made therefrom of preferably at least 20, more preferably at least about 25.

The figure of merit for a helium/methane separation is calculated by the equation V:

where PHe is the permeability of helium and αHe/CH4 is the selectivity of helium/methane. The values and derivatives are calculated at 30°C. The glassy polymers useful in the present invention have a figure of merit for helium/methane separation for membranes made therefrom of preferably at least about 5, more preferably at least about 15.

The membranes useful in the present invention preferably have an oxygen/nitrogen selectivity of at least about 10, more preferably at least about 12, and preferably an oxygen permeability of at least about 1 Barrer at a temperature of 5°C or less, as measured using a mixture of 80 mole percent nitrogen and 20 mole percent oxygen as a feed at a pressure of 30 psia on the first side of the membrane and a vacuum of less than 1 mm Hg on the second side of the membrane.

The membranes useful in the present invention preferably have a carbon dioxide/methane selectivity of at least about 40, more preferably at least about 80, and preferably a carbon dioxide permeability of at least about 10 Barrers at a temperature of 5°C or less, as measured using a mixture of 5 Mole percent carbon dioxide and 95 mole percent methane as feed with a pressure of 30 psia on the first side of the membrane and a vacuum of less than 1 mm Hg on the second side of the membrane.

The membranes useful in the present invention preferably have a helium/methane selectivity of at least about 70, more preferably at least about 100, and preferably a helium permeability of at least about 10 Barrers at a temperature of 5°C or less, as measured using a mixture of 5 mole percent helium and 95 mole percent methane as the feed with a pressure of 205.5 kPa (30 psia) on the first side of the membrane and a vacuum of less than 0.133 kPa (1 mm Hg) on the second side of the membrane.

Several glassy polymers are described on the following pages. However, it should be emphasized that only glassy polymer compositions defined in the lower section of page 6a and in the appended claims are part of the present invention.

The glassy polymers used herein are preferably those which, with respect to a gas mixture of interest, exhibit large increases in a separation factor with moderate decreases in a gas permeability of the more permeable component gas. Illustrative glassy polymers include certain polyphenylene oxides, certain polyimides, certain polycarbonates, certain polyamides, certain polyethers, certain polyester carbonates, certain polyarylates, certain polyesters, certain polyacetylenes, certain polytriazoles, certain polyoxadiazoles, certain polyolefins, certain polysulfones, certain polyether sulfones, certain polyamide imides, certain te polyvinylsilanes and certain polydenzoazoles such as polybenzoazoles, polybenzothiazoles, polybenzoimidazoles, polybenzobisoxazoles, polybenzobisthiazoles and polybenzobisimidazoles, and copolymers and physical mixtures thereof. For oxygen/nitrogen separation, preferred glassy polymers include certain polycarbonates, certain polyestercarbonates, certain polyarylates, certain polyimides, certain polyesters, certain polyethers, certain polyphenylene oxides, certain polytriazoles, certain polyoxadiazoles, certain polyolefins, certain polysulfones, certain polyethersulfones, certain polybenzoazoles, and copolymers and physical mixtures thereof. For carbon dioxide and light hydrocarbon separation, preferred glassy polymers include certain polybenzoazoles, certain polycarbonates, certain polyesters, certain polyestercarbonates, certain polysulfones, certain polyestersulfones, certain polyimides and copolymers and physical mixtures thereof. In general, the most preferred organic polymers useful in this invention generally have a rigid structure, such as a polybenzoazole or a polymer consisting of moieties of 9,9-bis(3,5-dibromo-4-hydroxyphenyl)fluorene, 9,9-bis(3,5-dichloro-4-hydroxyphenyl)fluorene, hexafluorobisphenol A, tetrahalohexafluorobisphenol A or tetraalkylhexafluorobisphenol A selected from the group consisting of polyimide, polyether, polysulfone, polyethersulfone, polyarylate, polyester, polyester carbonate, polycarbonate and copolymers and physical mixtures thereof. Hexafluorobisphenol A as used herein refers to bisphenol A in which all of the hydrogens of the isopropylidene bridging group have been replaced by fluorine moieties. Tetrahalo or tetraalkyl as used herein with respect to bisphenol A refers to bisphenol A in which four of the hydrogens of the aromatic rings have been replaced by halogen or alkyl moieties, respectively. Certain inorganic poly mers should also be usable. For example, carbon membranes, aluminum membranes and silicone-containing or other ceramic membranes can have the desired properties.

Other glassy polymers include substituted polycarbonates such as tetrachlorohexafluoro bisphenol A polycarbonate (TCHF BA PC) or tetrabromo-hexafluoro bisphenol A polycarbonate (TBHF BA PC), which have a repeating unit of the following formula.

where X is either Cl or Br. The polycarbonate may be a copolymer containing bisphenols in addition to tetrahalohexafluorobisphenol A, as well as ester moieties. A synthesis of such substituted polycarbonates and related polyestercarbonates is described in U.S. Patent 4,874,401. See also U.S. Patents 4,851,014 and 5,000,763.

Preferred glassy polymers according to the invention include polycarbonates containing 9,9-bis(3,5-dibromo-4-hydroxyphenyl)fluorene (TBF PC) or 9,9-bis(3,5-dichloro-4-hydroxyphenyl)fluorene (TCF PC), which have a repeating unit of the following formula:

where U is either Br or Cl. The polycarbonate may be a copolymer containing bisphenols in addition to 9,9-bis(3,5- dibromo-4-hydroxyphenyl)fluorene or 9,9-bis(3,5-dichloro-4-hydroxyphenyl)fluorene, as well as ester moieties.

Such fluorene-containing polycarbonates can be obtained by reacting the fluorene-containing bisphenol with a phosgene or other carbonate precursor using either solution or interfacial polycarbonate synthesis techniques. Examples of syntheses of polycarbonates according to these techniques are set forth in U.S. Patents 3,248,414; 3,153,008; 3,215,668; 3,187,065; 3,028,365; 2,999,846; 2,999,835; 2,964,974; 2,970,137; and 1,991,273. See also H. Schnell, Chemistry and Physics of Polycarbonates, Interscience Publishers, New York, 1964, pages 9-61.

Such fluorene-containing polycarbonates can be cast from a solution to form films or extruded from blends to form films and hollow fibers. The solutions or blends contain the polycarbonate, a solvent, and optionally a non-solvent. Preferred solvents for such fluorene-containing polycarbonates are include methylene chloride, chloroform, trichlorethylene, ortho-dichlorobenzene, N,N-dimethylacetamide, N,N-dimethylformamide, tetrahydrofuran, cyclohexanone 1,2,3,4-tetrahydronaphthalene, 1-formylpiperidine, 1-acetonapthone, 1-acetylpiperidine, benzaldehyde, 1-chloronaphthalene, cyclopentanone, Dimethyl-propionamide, ethyl-1-pyrrolidinone-2,1-cyclobenzyl-2-pyrrolidinone, 1,1,3,3-tetramethylurea, isophorone, 1,2,4-trichlorobenzene, N-formylmorpholine, N-methyl salicylate, N-methylcaprolactam, methyl benzoate, ethyl benzoate, diglym, chlorobenzene, 4-ethyl morpholine, methoxy-1-naphthalene, Phenetol, pyridine, pyrrolidinone-2, styrene, styrene oxide, N-methyl-1-pyrrolidinone, acetphenone, anisole, and 1,4-dioxane. Preferred nonsolvents for such fluorene-containing polycarbonates include acetone, acetonitrile, bis(2-ethoxyethyl)ether, diethylene glycol dibutyl ether, butyl stearate, n-butyl acetate, carbon tetrachloride, cyclohexane, decahydronaphthalene, decyl alcohol, diacetone alcohol, diethylene glycol, dimethyl carbonate, dimethyl malonate, dimethyl sulfone, dioctyl phthalate, dipropyl carbonate, dodecane, ethanol, ethyl acetate, ethyl formate, ethylbenzene, ethylene carbonate, ethylene glycol, ethylene glycol diethyl ether, ethylene glycol dimethyl ether, 1-hexanol, isopropyl alcohol, methanol, methyl acetate, methylene caproate, methyl caprylate, methylene myristate, methyl valerate, methylcyclohexane, nitromethane, polydimethylsiloxane, polyethylene glycol E600, propionitrile, 1,1,1- Trichloroethane, triethylamine, triethylene glycol, tetraglyme, triglyme, diphenyl ether, polyethylene glycol 1500, propylene carbonate, N-methyl acetamide, and xylene. The membrane is formed by casting the solution or by extruding the mixture and evaporating or leaching the solution and optionally the non-solution. The membrane can also be coagulated or leached after casting or extrusion by bringing the membrane into contact with a suitable non-solvent such as water. Membrane properties can can be modified by subsequent stretching and/or subsequent heat treatment of the membrane.

Another group of glassy polymers are substituted polyesters, such as tetrabromohexafluoro bisphenol A isophthalate ester (TBHF BA IE) or tetrachlorohexafluoro bisphenol A isophthalate ester (TCHF BA IE), which have a repeating unit of the following formula:

where Y is either Cl or Br. The polyester may further be a polymer containing bisphenols in addition to tetrahalohexafluoro bisphenol A.

Certain polybenzoxazole (PBO) or polybenzothiazole (PBT) materials are particularly suitable for use as membranes. A PBO or PBT which is not a highly crystalline, rigid rod structure is preferred. For example, preferred polybenzoxazoles include those which comprise repeating units of the following formula: PBO A

or PBO B

where Z is oxygen.

In preferred polybenzothiazoles, Z is sulfur in the aforementioned repeating units. Isomers of the repeating units of PBO A and PBO B are also possible. Furthermore, the following repeating units can be used:

or.

The PBO and other polybenzoazole materials can be prepared as described in U.S. Patents 4,847,350; 4,772,668; and 4,533,693 and Japanese Kokai 62/207,322 and 01/159,024. U.S. Patent Application Serial No. 513,345, filed April 20, 1990, describes the preparation of microporous PBO membranes. U.S. Patent 5,034,026, published July 23, 1991, describes PBO membranes for gas separation. U.S. Patents 4,939,235; 4,963,428; 4 487 735 and 4 898 924 as well as Japanese Kokai 61/28 530 describe the preparation of films from polybenzoazoles. The polybenzoazoles are generally soluble in polyphosphoric acid, but certain PBO and PBT materials are more easily formed as gas separation membranes using more volatile solvents such as methanesulfonic acid. Additional groups can be incorporated into the PBO or PBT introduced to make the polymer more amenable to the processing steps and operations suitable for forming the membrane. The PBO or PBT, particularly an indane-containing PBO such as PBO A, can be cast or extruded from a solution of methanesulfonic acid to form films. For an indane-containing PBO, m-cresol and certain ether-organic solvents can dissolve the polymer. If m-cresol is the solvent, the solvent can be evaporated from the membrane or washed out with a non-solvent. The film can also be coagulated after casting by contacting it with a suitable non-solvent. Membrane properties can be modified by stretching and/or preferably by heat treating the membrane. Heat treatment is desirable in that it eliminates surface pores in the membrane. Generally, the membrane must be carefully dried before use, that is, the solvent and coagulants must be removed, in order to achieve the desired properties.

In one embodiment, an asymmetric PBO membrane can be made from a dope containing a solvent for the polymer and a non-solvent pore-forming compound. Advantageously, the dope is coagulated in a polymer precipitant which leaches out a portion of the solvent as well as a portion of the pore-forming compound. For the preparation of polyphenylene oxide membranes, see U.S. Patents 3,350,844; 3,709,774; 3,852,388; 4,468,500; 4,468,501; 4,468,502; 4,468,503; and 4,949,775. U.S. Patent 4,971,695 describes sulfonated and/or hexafluorobisphenol A-containing polysulfone membranes.

US Patent 3,899,309 describes the synthesis and preparation of aromatic polyimide, polyester and polyamide release membranes. U.S. Patents 4,717,393; 4,717,394; 5,009,679; 5,034,027; 5,042,992; and 5,042,993 describe certain polyimide membranes. U.S. Patents 5,007,945 and 5,013,332 describe certain polyarylate membranes. While these membranes were not evaluated in the process of the present invention, useful membranes can be readily identified by experimentation.

Certain polyimide materials are particularly preferred for use as membranes. Such polyimides include those containing indane moieties. Indane-containing, as used herein, refers to moieties based on a cycloaliphatic ring and two aromatic rings, one of which is joined to a five-membered ring which is saturated, except for the two carbon atoms which are part of the aromatic group, while the remaining aromatic ring is attached by a single bond to the saturated five-membered ring, which are connected to the polymer by at least one single bond from each aromatic ring. Indane-containing polyimide, as used herein, refers to a polyimide which contains at least a portion of diamine residues derived from indane-containing diamines. The polyimides can be derived from aliphatic, alicyclic and/or aromatic dianhydrides and aliphatic, alicyclic and/or aromatic diamines, provided that at least a portion of the diamines contains indane moieties. Such polyimides include polyimides derived from mixed dianhydrides and/or mixed diamines, provided that at least a portion of the diamines contains indane moieties. Preferably at least 25 percent, more preferably at least 50 percent, even more preferably at least 75 percent, most preferably 100 percent of the diamine moieties present in the polyimide contain indane moieties.

The indane-containing polyimides preferably comprise repeating units corresponding to the following formula:

where

R is independently selected at each occurrence from the group consisting of:

A. a tetravalent phenylene, naphthalene or perylene radical;

B. a tetravalent bisphenylene residue according to the following formula:

where

L is independently selected at each occurrence from the group consisting of a single bond, -O-, -CO-, -S-, -SO-, -SO₂-, -SS-, a divalent C₁₋₁₂ hydrocarbyl group, and a divalent C₁₋₆ halohydrocarbyl group; and

C. a tetravalent bisphenylene radical according to the following formula:

where

W is independently selected at each occurrence from the group consisting of hydrogen, a monovalent C₁₋₈ hydrocarbyl radical, and a monovalent halohydrocarbyl radical; and

R¹ is independently at each occurrence a divalent radical which corresponds to the following formula:

or.

where

V at each occurrence is independently selected from the group consisting of hydrogen, a monovalent

C₁₋₆ hydrocarbyl or halohydrocarbyl radical, a monovalent C₂₋₆ oxyhydrocarbyl or oxyhalohydrocarbyl radical, -OH, - Cl, -Br, -NO₂ and -SO₃H; and

W' at each occurrence is independently selected from the group consisting of hydrogen, a monovalent C₁₋₈ hydrocarbyl group, and a monovalent C₁₋₈ halohydrocarbyl group.

Such indane-containing polyimides can be prepared by reacting an aliphatic, alicyclic and/or aromatic dianhydride with an indane-containing diamine in the presence of a base such as an isoquinoline and a solvent such as a m-cresol. The solution is sparged with an atmosphere of a noble gas such as dry nitrogen, refluxed for about 1 hour to 30 hours, cooled to about 120°C to about 50°C, precipitated in alcohol such as methanol, and recovered by filtration. The recovered polymer is optionally washed or redissolved and reprecipitated to further purify the polymer and preferably dried under vacuum at about 80°C to about 150°C. Such polyimides useful in the present invention can also be prepared by methods disclosed in the art. See U.S. Patents 4,366,304; 4,378,400; 4,454,310; 4,474,858; 4,639,485; 4,645,824; 4,687,611; 4,742,153; and 4,838,900.

Such indane-containing polyimides can be cast from solution to form films or extruded to form films and hollow fibers. See U.S. Patents 4,746,474; 4,880,699; 4,929,405; and 4,978,573. Preferred solvents for such polyimides include m-cresol and dipolar aprotic solvents such as N-methyl-2-pyrrolidinone, N,N-dimethylacetamide, and dimethylformamide. The solvent is evaporated from the membrane or washed from the membrane with a nonsolvent such as water. The membrane can also be coagulated after casting or after extrusion by contacting it with a nonsolvent. Membrane properties can be modified by stretching and/or heat treating the membrane.

Certain substituted polyethers are also said to be preferred materials for membranes used in the present invention, particularly polyethers containing hexafluoro-bisphenol-A moieties. The synthesis of such polyethers is known in the art. See Mercer et al., International Publication Number WO 91/09081. Membranes can be formed from such polymers by casting a film of a solution of the polymer in a suitable solvent, such as a chlorinated hydrocarbon.

Certain substituted polytriazoles and polyoxadiazoles are also preferred materials for use in the present invention. Polytriazoles and polyoxadiazoles comprising substituted alkyl, aryl and halogen-containing moieties can be fused by condensation polymerization and formed into membranes by casting or extruding suitable solutions or mixtures containing the polymer. For example, the preparation of porous or liquid separation membranes from oxadiazole polymer is discussed in Japanese Patent No. 79/25,278; U.S.S.R. Patent Nos. 1,033,510; 1,038,347; and 1,248,629; and WO 8200648. Gas separation membranes made of such polymers can be cast or extruded from solutions or mixtures of the polymer in a suitable solvent such as sulfuric acid, methanesulfonic acid, trifluoroacetic acid, formic acid, polyphosphoric acid and the like. For additional information regarding polytriazole membranes see Gebben et al., "Gas Separation Properties of a Thermally Stable and Chemically Resistant Polytriazole Membrane", Journal of Membrane Science, Vol. 46, pages 29-41 (1998).

In general, polymers useful for gas separation membranes exhibit relatively consistent and reproducible transport properties in terms of gas permeability and selectivity, which are independent of the membrane manufacturing process. For such polymers, the resulting membrane performance depends essentially only on the chemical composition of the membrane and its effective thickness. However, some polymers exhibit variable transport performance depending on the membrane manufacturing methodology and membrane history. More specifically, polymers that exhibit freezing points well above ambient temperature are sensitive to the membrane manufacturing process and membrane history. This behavior is believed to be due to the thermodynamic non-equilibrium state of these polymers. The amount and distribution of free or unoccupied volume and the degree of orientation and/or crystallinity can influence the gas transport properties of such polymers.

Such behavior has been observed in the art. For example, U.S. Patents 4,880,441 and 4,871,494 disclose methods for influencing the amount of free volume in a membrane separation layer by changing the manufacturing process. High concentrations of carbon dioxide dissolved in a membrane can be used to change the free volume state of the separation layer of the membrane, as described in U.S. Patent 4,755,192. Procedures as described in these patents provide methods for changing the permeability and selectivity properties of a membrane material, either during manufacture of the membrane or as a subsequent treatment. These methods can be useful for making membranes that meet the criteria of the present invention. In some cases, a material that outside the scope of the present invention by using a particular membrane manufacturing process can be made to fall within the scope of the present invention by changing the manufacturing process. It is also possible to use a process after completion of the manufacture of the membrane, such as swelling the membrane by exposing it to carbon dioxide or another gas or liquid, which changes the original performance of the membrane so that it falls within the scope of the present invention. In addition, other membrane modification processes such as surface modification, UV irradiation, cross-linking, monoaxial or biaxial stretching or addition of additives, which are known to those skilled in the art, can be used to produce membranes that meet the criteria of the present invention.

Therefore, a polymeric material that can be made into a membrane that falls outside the scope of the present invention should not necessarily be considered outside the scope of the present invention in every situation, since a number of manufacturing or membrane modification processes are available for changing the performance of the polymer in the final membrane form. These processes can be used to produce membranes that have the properties desired for the present invention.

Fig. 12 shows that only certain gas separation membranes meet the criteria set forth for the present invention. Fig. 12 shows actual performance data, represented by open circles for a first membrane (Membrane 1) and by open squares for a second membrane (Membrane 2). This data was generated by performing of separations at different temperatures with respect to oxygen and nitrogen. Curves were generated by applying Arrhenius interpolations and Arrhenius extrapolations from actual data on the performance of two different membranes with respect to the separation of oxygen and nitrogen at different temperatures. The filled circles and squares indicate an interpolated and an extrapolated performance at temperature intervals of 20°C for the two different membranes. The curve for membrane 1 shows a membrane performance curve fitted to actual data with respect to a TCHF BA PC membrane. The curve for membrane 2 shows performance of cellulose triacetate, a commercially available membrane material, which is presented as a comparative experiment which does not fall within the scope of the invention. It can be seen that at -20°C the first membrane achieves a separation factor of about 7.6 and an oxygen permeability of about 4.6 Barrers. The second membrane, which is not used in the process of the present invention, has a separation factor of about 9.2 at -20°C and an oxygen permeability of about 0.4 Barrers. Many prior art membranes, as represented by membrane 2 in Figure 12, generally have either too low a gas permeability at reduced temperatures or relatively little improvement in selectivity. Other examples of membranes which do not have the desired properties are membranes made of TRYCITE polystyrene, tetramethyl bisphenol A polycarbonate, unsubstituted polysulfone and tetrabromo bisphenol A polycarbonate. The membranes useful in the present invention can be in the form of flat films, hollow fibers or hollow tubes. The morphology or structure of the membrane can be homogeneous, composite or asymmetric.

The structure or morphology of membranes possible in the process of the present invention can be homogeneous. In the case of a homogeneous membrane, the entire thickness of the membrane can serve as a separation layer. Homogeneous membranes are made by forming a continuous, thin separation layer which is dense and substantially free of voids or pores. In general, homogeneous membranes have a relatively low flux through the membrane due to the thickness of the separation layer of the membrane.

The entire membrane need not be made of the glassy polymer. In many embodiments, it is preferable to use a composite membrane with a porous substrate on which the glassy polymer separating layer is formed or deposited. Such porous substrates may be isotropic or anisotropic. A composite membrane may be multilayered and may have several support, intermediate and collection layers in addition to the separating layer. The substrate may be any suitable material having the required strength and permeability. Polysulfones, polyethersulfones, aromatic polyimides, polycarbonates, polyester carbonates, polyesters, polyether ketones, polyamides, polyolefins, polytetrafluoroethylenes and polybenzo-azoles are preferred as porous substrate materials. Porous inorganic supports such as porous ceramics are also possible. The separation layer can be cast on the substrate, plasma polymerized on the substrate, laminated, or formed by any other suitable technique, so long as the separation layer is substantially defect-free, that is, substantially free of pinholes or other discontinuities that adversely affect membrane performance. It is also desirable that the separation layer be as thin as possible to maximize the flux of gas permeating through the membrane.

Asymmetric or anisotropic membranes may also be used. Such membranes have at least one dense separation region and at least one generally porous or less dense region which represents an additional mechanical support and is formed from the same polymer. In the embodiments in which the membrane is a hollow fiber, the separation region can occur at or near the external outer surface, at or near the internal inner surface, at a region between the external and internal surfaces, or at a combination thereof. In one embodiment, the hollow fiber membrane comprises a separation region at or near the inner surface of the hollow fiber membrane and a generally porous region which extends from the separation region to the outer surface of the hollow fiber membrane. In another embodiment, the hollow fiber membrane comprises a generally porous region at or near the outer surface of the hollow fiber membrane, a generally porous region at or near the inner surface of the hollow fiber membrane, and a separation region which is arranged generally between the two porous surface regions. The separation region in the hollow fiber membranes of the present invention functions to selectively separate at least one component gas from at least one other component gas in the feed gas mixture. The separation region in such membranes may be a dense region, a region of non-continuous porosity, a region resembling a closed foam cell, or a region having an enhanced free volume state. Anisotropic or asymmetric membranes are suitably formed by casting or extruding the polymer into mixtures of a solvent and an optional non-solvent with respect to the polymer generally described in the prior art. Illustrative patents which describe the manufacture Nos. 4,955,993; 4,772,392; 4,486,202; and 4,329,157. The choice of solvent and preferred concentration of polymer will vary depending on the solvent, polymer, manufacturing process, and other factors. Subsequent heat or stretching treatments may be used to modify the properties of the membrane. Optionally, the membranes may then be subjected to other treatments described in the art, including surface treatment by processes such as solvent annealing, etching, irradiation, cross-linking, fluorination, sulfonation, plasma treatment, and the like.

Flat film membranes useful in the present invention have a thickness between about 25 micrometers and about 500 micrometers. Hollow fiber membranes useful in the present invention have a diameter in the range of about 40 micrometers to about 750 micrometers, preferably in the range of about 75 micrometers to about 500 micrometers. The ratio of outside diameter to inside diameter of such hollow fiber membranes is preferably between about 1.15 and about 2.50, more preferably between about 1.25 and 1.70. In the case of asymmetric or composite membranes, the separation layer is preferably less than about 10 micrometers, more preferably less than about 1 micrometer, even more preferably less than about 0.5 micrometers. The separation layer in asymmetric or composite membranes can be located on the outer surface or the inner surface of the hollow fiber membrane.

The membrane can be incorporated into a variety of designs. Flat foil or flat film membranes can be either in a plate and frame or in a spiral winding Hollow fiber and tubular devices may also be used. In the case of hollow fiber devices, the mixture to be separated may be introduced into the bores of the hollow fibers or may be located externally of the hollow fibers. Generally, such hollow fiber membranes are formed into hollow fiber bundles as described in U.S. Patents 3,228,876; 3,422,008; and 4,881,955. The hollow fibers may be prestressed or spiraled around a central core or arranged parallel in the hollow fiber bundle with or without a core. The core may be perforated, allowing the core to be used to collect or distribute gas externally of the hollow fibers. The core may be flexible or collapsible to accommodate changes in fiber length with temperature. See, for example, U.S. Patent 5,026,479. Preferably, each of the hollow fibers has a wall thickness, outside and inside diameters, lengths, interface thickness, and other properties within a narrow range, since a significant change in fiber properties can reduce overall performance. Changes of less than 5 percent in such properties are preferred. Typically, a tubesheet is present at each end of the hollow fiber bundle, but other designs, such as a single center tubesheet, are possible. The tubesheet can be made of any resin that can maintain the desired mechanical strength at operating temperatures. Epoxy resins, polyurethane resins, or elastomers are generally preferred as tubesheet materials.

If the membrane module is subjected to repeated cooling and heating cycles, provisions must be made for the different rates of thermal expansion of the component parts. Alternatively, the materials used in the construction of the module can be selected so that that the dimensional changes during cooling and heating do not unduly stress and damage the membrane elements. Compressive and tensile forces or stresses on the components of a membrane device can be generated during low temperature operation and/or low/high temperature cycling. The cause of such forces and stresses can be the large difference in the coefficients of linear thermal expansion (CLTE) of the various materials used in the manufacture of a membrane device. Traditionally, membrane devices have been manufactured using metals for the vessel or housing or optional core tube, while thermosets such as epoxies have been used for the tube sheet and thermoplastic materials have typically been used for the membrane. The seals used in such devices are usually butadiene elastomers such as nitrile rubber. The CLTE values of most materials are typically less than 2.5 x 10-5 cm/cmºC. The rubber, thermoset or thermoplastic materials used in membrane devices typically have CLTE values in the range of 3.0 to 10.0 x 10-5 cm/cmºC.

If there are large differences in CLTE values between component parts, low temperature operation can cause circumferential cracking at the core tube/tubesheet interface as well as at the membrane/tubesheet interface. In extreme cases, the membranes can actually separate from the tubesheet because the thermoplastic membranes shrink faster than the metal core tube. Using a plastic core tube with a CLTE value similar to the CLTE value of thermoplastic membranes or using a hinged core tube significantly reduces or eliminates this problem. Alternatively, making the membrane length longer than the distance between the tubesheets allows the membrane to shrink without creating stress.

Typically, the seal between the membrane module and the vessel or housing is an O-ring or gasket formed from an elastomeric material such as nitrile rubber. The operating limit for most butadiene-based elastomers is -40ºC. Below -40ºC, seals made from alternative materials should preferably be used, such as O-rings or gaskets made from silicone rubber, derivatives of silicone rubber such as phenyl substituted silicone rubber, or metal such as aluminum or an aluminum alloy. Alternative designs such as an integrated vessel/tube sheet, which would eliminate the need for O-rings or gaskets, can also be used.

Cooling the device from ambient temperature to the low operating temperature can cause leakage between the vessel and the tubesheet because the tubesheet shrinks faster than the vessel. Numerous solutions to this problem are possible, such as proper sizing of the O-ring or gasket, or using a vessel made of a material with a closer match between CLTE and the tubesheet. In some cases it may be necessary to modify the standard materials used to make the tubesheet, i.e. epoxy or urethane resins. Blends with materials more suitable for low temperature operation may be necessary. Silicone rubber or silicone rubber derivatives can be blended with the epoxy or urethane. In extreme cases the standard epoxy or urethane tubesheet can be completely replaced with improved low temperature materials such as silicone elastomers or silicone derivative elastomers.

In one embodiment, the feed gas mixture to be separated is introduced into the holes of hollow fibers and gas permeating through the fiber walls generally flows in a substantially countercurrent direction with respect to the gas mixture in the fibers. U.S. Patents 4,871,379; 4,929,259; 4,961,760; and 5,013,437 describe illustrative devices. See also U.S. Patent Application Serial No. 07/769,040, filed September 30, 1991. Countercurrent flow of the permeating gas can be promoted by concentric non-permeable jackets in the bundle. Such jackets can extend from the tubesheet through which the non-permeate gas exits the fiber in the bores of the hollow fibers. If the end of the jackets is embedded in the tubesheet, the jacket material should be selected so that there is good adhesion between the jacket and the tubesheet at low operating temperatures; otherwise, delamination may occur. With respect to epoxy tubesheet materials, suitable cladding materials may include polyvinyl chloride and acetate films. The claddings preferably terminate near, but not in contact with, the opposite tubesheet through which feed gas is introduced. Gas permeating through the walls of the hollow fibers may enter a hollow fiber or core in the bundle through a hole or holes in the tube near the tubesheet through which feed gas is introduced.

In some embodiments, it may be desirable to use a membrane device adapted to receive a purge fluid. See U.S. Patents 4,961,760; 5,013,437 and 5,026,479 as examples of such devices. Such membrane devices have an inlet for the purge fluid, which is preferably a gas. In some embodiments, part or all of the permeate gas or the non-permeate gas may be used as the purge gas. The purge fluid preferably flows in a Direction which is substantially opposite to the flow direction of the feed gas mixture within the device.

Single or multiple membrane elements can be arranged in series, parallel, cascade or in recycle configuration. Multiple membrane devices can be arranged in parallel so that the feed gas mixture is split and a portion of the feed gas mixture is sent to each membrane device. Permeate and non-permeate streams from each device can be combined with similar streams from the other devices. Multiple membrane devices can also be arranged in series so that the higher pressure non-permeate stream from one device provides the higher pressure feed stream to the other device. Alternatively, multiple membrane devices can be arranged in cascade so that the lower pressure permeate stream from one device is the higher pressure feed stream to another device. Single or multiple devices can be arranged for recycle operation so that a portion of the permeate or non-permeate stream from one device can be recycled to be used as a portion of a feed stream to the same device or another device. For example, ST Hwang and JM Thorman in AIChE Journal, Vol. 26, No. 4, pages 558-566 (1980) describe the use of multiple membrane devices arranged in a continuous membrane column arrangement. SA Stern, JE Perrin and EJ Naimon in Journal of Membrane Science, Vol. 20, pages 25-43 (1984) describe the use of recycle and multiple permeator systems. JE Perrin and SA Stern in AIChE Journal, Vol. 32, pages 1889-1901 (1986) describe the use of membrane systems using two different membrane materials to improve separation efficiency. It is also possible to use a variety of arrangements within a membrane module. For example, M. Sidhoum, S. Majumdar and KK Sirkar in AIChE Journal, Vol. 35, pages 764-774 (1989) describe an internally staged hollow fiber module which uses two membranes in the same device. Other arrangements of membrane devices as well as combinations of the foregoing arrangements of multiple membrane devices are possible. The membrane material or composition used in one element, device or stage may be different from the membrane material or composition used in another element, device or stage. U.S. Patent 4,894,068 describes a system which uses membrane elements in series. A single membrane element can be divided by baffles to provide a plurality of separation sections which can be arranged in series to achieve higher purity or in parallel to obtain variable flow. Examples of such devices can be found in U.S. Patents 5,013,437 and 5,013,331.

The preferred glassy polymer membranes do not necessarily exhibit selective separations of all components of gas mixtures. To simplify the description of component gases to be separated, it is convenient to consider pairs of gases. Minor components and impurities that do not deleteriously affect the physical integrity or performance of the membrane may also be present in such gas mixtures, but for simplicity, the major component gas pairs will be discussed below. The membranes are useful for separating preferably at least one of the following component gases from a gas mixture: hydrogen, helium, oxygen, nitrogen, argon, Carbon monoxide, carbon dioxide, steam ammonia, hydrogen sulfide, nitrogen oxides, sulfur oxides or light hydrocarbons. Light hydrocarbons as used herein refer to gaseous saturated and unsaturated C1-5 hydrocarbons such as methane, ethane, ethene, ethyne, propane, propene, propyne, propadiene, n-butane, 2-methylpropane, 2-butene, 1-butene, 2-butyne, 1-butyne, 1,3-butadiene, 1,2-butadiene, n-pentane, cyclopentane, 2-methyl-2-butene, 3-methyl-1-butene, 2-methyl-1-butene, 2-pentene, 1-pentene, 2-methyl-1,3-butadiene, 3-methyl-1,2-butadiene, 2,3-pentadiene, 1,4-pentadiene, 1,3-pentadiene, 1,2-pentadiene, 3-methyl-1-butyne, 2-pentyne, 1-pentyne, and 2-methylpropane.

It is believed that the glassy polymer membrane is most effective at separating gas pairs when one gas of the pair has both higher solubility in the glassy polymer and higher diffusivity through the glassy polymer. This is the case with mixtures of oxygen and nitrogen and carbon dioxide and light hydrocarbon. Other gas pairs are possible even if they do not have this desired relationship. For example, hydrogen or helium and light hydrocarbon, which can be advantageously separated by the process of the present invention, do not have this desired relationship between solubility and diffusivity. In general, the more soluble component gases are more polarizable or polar, which can be demonstrated by boiling point, critical temperature measurements or other conventional techniques. Diffusivity is related to the size of the gas molecule, with smaller gas molecules generally diffusing faster than larger gas molecules. Preferred gas pairs which can be easily separated using glassy polymer membranes at low temperatures include: (1) oxygen and nitrogen, as in Air is present, (2) carbon dioxide, optionally with other acidic gases such as H₂S, SOx or NOx, and light hydrocarbon, (3) oxygen and argon, nitrogen and methane, and (5) hydrogen and methane. Separation of mixtures of oxygen and nitrogen and mixtures of carbon dioxide and a light hydrocarbon is particularly preferred.

Other gas pairs which have not been fully evaluated but are believed to be suitable for separation by the process of the present invention include: (1) hydrogen or helium and one or more gases selected from nitrogen, ammonia, carbon dioxide and a light hydrocarbon; (2) nitrogen and one or more gases selected from argon, methane or a mixture of light hydrocarbons; (3) hydrogen chloride and chlorine; (4) methane and C2 or C3 alkanes or alkenes, respectively; and (5) ethane and ethylene.

The feed gas mixture should advantageously be free of contaminants that may adversely affect the physical integrity or performance of the membrane, other parts of the membrane device, or the system equipment. For example, certain contaminants may adversely affect membrane performance or system components over time. Contaminants present in the feed gas stream to a membrane device can cause a variety of performance degradation effects. For example, H. Finken in Material Science of Synthetic Membranes, ACS Symposium Series No. 269, American Chemical Society, Washington, DC, page 229 (1984), describes the effects of high concentrations of carbon dioxide dissolved in a membrane on membrane performance. E. Sanders in Journal of Membrane Science, Vol. 37, pages 63-80 (1988), provides additional information regarding Carbon dioxide exposure. These effects are generally referred to as transport plasticization. Typically, a membrane that has undergone transport plasticization will exhibit reduced selectivity and increased permeability to one or more gases. At low temperature, the concentration of dissolved gas (which can cause transport plasticization in a membrane) is generally higher for a given gas phase partial pressure of that gas. This leads to increased susceptibility to transport plasticization effects at low temperatures. These effects may be caused by contaminants or impurities in the feed gas stream or, in some cases, by one of the primary stream component gases to be separated.

An additional class of fouling problems arises from pore-clogging effects or competitive sorption and diffusion effects. These effects lead to a decrease in the permeability of the membrane. These effects are also generally greater at lower temperatures. For example, K. H. Lee and S. T. Hwang, Journal of Colloid and Interface Science, Vol. 110, pages 544-554 (1986) report examples of pore clogging by condensable vapors. These pores can be located in the porous support layer of a membrane. This mechanism can also operate in the larger free volume elements in a glassy polymer. This effect typically produces a decrease in the permeability through the membrane.

It is therefore important to provide adequate pretreatment of gas feed streams to membranes operating at low temperatures. This pretreatment should preferably include removal of impurities or Impurities that adversely affect the selectivity or gas permeability of the membrane.

It is believed that certain hydrocarbons, compressor oils, glycols, oil field additives, surfactants, mercury, and solid particles or liquid droplets larger than 1 micron can foul some membranes used in the separation of gases. Water or other condensables can freeze and clog passages and spaces or reduce the effective membrane area or impair heat exchange capabilities. It may be necessary or desirable in some applications to remove hydrogen sulfide, hydrogen chloride, nitrogen oxides, sulfur oxides, chlorine, mercaptans, carbonyl sulfide, and other acidic or corrosive contaminants from the feed gas mixture before the mixture contacts the membrane device. The components of the gas mixture should not significantly plasticize or swell the glassy polymer to a detrimental degree. Consequently, certain gas components, such as carbon dioxide, should not be present in amounts that would excessively plasticize or otherwise have a detrimental effect on the membrane. Different membranes tolerate plasticizers and contaminants more or less well.

Therefore, the method and apparatus optionally include means for reducing contaminants, particularly those contaminants having a detrimental effect on the physical integrity or performance of the membrane, from the feed gas stream to the process. Such contaminants can be classified as corrosives, particulates, or condensables. Corrosives are compounds that damage the process equipment by reacting with it. Some equipment can be made from a corrosive sion resistant material. However, parts of equipment must be made of certain materials and corrosives which attack these materials should be removed. Such corrosive contaminants may include carbon dioxide, sulfur oxides, nitrogen oxides, hydrogen sulfide, chlorine, hydrogen chloride, acetylene, mercaptans, carbonyl sulfide and the like. Particles or fines which are entrained in a gas stream and can clog heat exchanger passages, clog fibers in a hollow fiber membrane device or pores in the membrane material. Particles can also cause mechanical erosion or abrasion of process equipment having moving parts such as compressors, expansion devices and fans. Preferably, particles larger than 1 micron in size are removed by pretreatment. Condensables are compounds which can condense or freeze from a gas stream at the desired operating temperature. Condensables also greatly reduce heat exchanger heat transfer. Processes may be able to tolerate a certain amount of liquid condensation, particularly if the equipment is oriented and outlets are provided so that such liquids can drain from the equipment. However, solid condensed compounds cause the same problems as particles and tend to accumulate more quickly. Such condensables may include water vapor, carbon dioxide hydrates, hydrocarbons and any other components present in an amount exceeding the vapor pressure thereof at any point in the process.

Impurities can be removed from the process by any conventional means. Condensable Substances may be removed by condensation in a specially designed process device. For example, water vapor can be condensed from air by cooling the air and then passing the cooled air into a separation vessel. Another device can remove the condensables while cooling a stream; recuperative heat exchangers used in cryogenic air separation plants cool air while allowing carbon dioxide and water to freeze on the walls, and process streams periodically switch passes in the exchangers to allow the accumulated solids to evaporate. Physical or chemical absorption is also commonly used to remove unwanted contaminants from a gas stream. Absorption in beds of molecular sieves of activated carbon, activated alumina, or zeolite is commonly used to remove contaminants. Physical filtration is often used to remove particles larger than 1 micrometer. Contaminants of a gas stream, such as water vapor and other condensables, may also be removed by permeation through a membrane device. Contaminants may be extracted by contact with a liquid stream which preferentially absorbs or reacts with the contaminant. Other known separation techniques as well as combinations of any of the foregoing techniques may be used.

In separating the component gases present in the gas mixture, the gas mixture is brought into contact with the first side of the membrane, this being done under the condition that a chemical potential difference exists across the membrane, whereby permeation through the membrane from the first side of the feed gas mixture of the membrane at a first chemical potential to an opposite second non-feed side of the membrane at a second chemical potential which is lower than the first chemical potential. Suitably, a partial pressure difference is used to provide the required chemical potential difference. The partial pressure of the feed gas components is suitably higher than that of the permeate gas components, resulting in a total transmembrane pressure in the range of from about 5 to about 2000 pounds per square inch (psi) (about 34.5 to about 13,800 kPa), and preferably from about 50 to about 1,200 psi (about 345 to about 8,270 kPa). The gas which does not permeate through the membrane, that is, the non-permeate gas, is generally at a slightly lower pressure than the feed gas mixture due to diffusion of a portion of the component gases through the membrane and the accompanying drop in pressure as the gas passes through the membrane device. Preferably, this pressure drop is kept to a minimum. The partial pressure difference can be provided by a feed gas stream available at high pressure, a feed stream compressor or blower, a permeate stream vacuum pump, an interstream downstream compressor, another type of compression device, or combinations thereof. A temperature or concentration difference can also be used to produce the required chemical potential difference across the membrane.

In some applications, depending on the gases to be separated, the membrane used, the volume of gas to be processed, the desired purity of the product and other factors, the separation preferably takes place at a temperature between about -30ºC and about 5ºC. In other embodiments, the gas in contact with the membrane preferably has a temperature of about -5ºC or less, and more preferably about -25°C or less. The temperature may vary over a wide range depending on the components of the feed gas mixture and the properties of the membrane used at the desired product purity. Advantageously, a temperature is chosen which achieves a high separation factor without excessive losses in permeability. The temperature must also be chosen such that the physical properties of the membrane or membrane device are not deleteriously affected. The selection of an operating temperature may also be limited by the energy required to adequately cool the feed gas mixture to a lower temperature and other pragmatic considerations.

If the gas mixture is not supplied at the desired temperature, any conventional means or combination of means for changing the temperature of the gas mixture, such as an external absorption or evaporative cooling system, an expansion device, or the self-cooling capability of the permeated gas due to the Joule-Thomson effect, may be used to achieve the desired temperature. The external cooling system as used herein refers to the use of a separate process or device for cooling one or more process streams that are in contact with the membrane by indirect heat exchange, thereby cooling the process. The indirect heat exchange as used herein refers to bringing two or more fluid streams into heat exchange relationship without any physical contact or mixing of the fluids. In one embodiment, a cold liquid coolant stream may be used to cool one of the process streams in a heat exchanger. The cold liquid coolant stream may be obtained the high pressure created by expansion of a gas stream. The coolant may be any fluid, such as freon gas, with favorable thermodynamic properties. Such external cooling systems are known and can be easily constructed by an expert in the field. The temperature of the gas mixture may also be changed by heat exchange with other process fluids, such as part or all of the feed gas, the non-permeate gas, or the permeate gas cooled by expansion, such as in a turbine expander.

Alternatively, the self-cooling ability of the permeate gas can be used for cooling. It is known that most gases have a significant cooling effect when the pressure is reduced without extracting energy from the gas. This phenomenon is known as the Joule-Thomson effect. The magnitude of this effect depends on the thermodynamic properties of the gas. For some gases, such as carbon dioxide at about -50ºC to 20ºC, the decrease in temperature associated with the pressure drop across the membrane can be sufficient to maintain a membrane device at a fairly low operating temperature. This effect is referred to below as the self-cooling properties of a gas. Membrane devices, such as those comprising hollow fibers, generally comprise a large membrane surface area to facilitate mass transfer between the streams in contact with the membrane. This large surface area also facilitates heat transfer between streams. Thus, membrane devices can advantageously be used as combined mass and heat transfer units. In one embodiment, as shown in Example 10, a gas mixture entering a hollow fiber membrane module cools due to the expansion of the permeating gas across the membrane. High heat transfer in a countercurrent flow arrangement results in cooling along the length of the membrane device, and the non-permeate product gas exits the device significantly cooler than the permeate product gas. In other embodiments, introduction of a purge stream having a different temperature than that of the feed gas stream can improve the mass transfer performance of the device and control the temperature of the non-permeate product gas exiting the device.

In certain embodiments, the feed gas may be available at a temperature that is lower than the desired operating temperature. In such embodiments, it may be necessary to heat the feed gas to the desired operating temperature. This may be accomplished using any conventional heating means, such as using an electric heater, a heat exchange with a warmer fluid, a heat exchange with a warmer ambient, and the like.

Generally, the preferred temperature for the feed gas mixture for the separation of oxygen and nitrogen is in the range of about -150ºC to about -5ºC, more preferably about -125ºC to about -25ºC, and most preferably about -100ºC to -30ºC. Generally, the preferred temperature for the separation of carbon dioxide and a light hydrocarbon is in the range of about -40ºC to 5ºC, or about 0ºC or higher.

During continuous operation, the temperature of particulate streams in the process is advantageously controlled to maintain a stable operating temperature for the membrane device and other process equipment, which is preferably within at least +/-5ºC. The temperature of the gas can be monitored by means of series-arranged thermocouples or other temperature monitoring devices. Conventional methods of temperature control are generally available to an expert in the art and are suitable for most processes. In processes using turbo-expanders, adjustable vanes provide some degree of control. Fine control in these processes can be obtained by passing a portion of the process stream through a control valve, bypassing either the turbo-expander or a heat exchanger.

The membrane device, piping and/or other devices of the device or system are preferably insulated using an insulating material such as expanded perlite to minimize energy losses and operating costs associated with cooling requirements in this process. Such provisions are particularly important where the system is not operating continuously or intermittently on demand, or where rapid start-up of the system is desired. In some embodiments, it may be desirable to initially pass more of the feed gas mixture through the turbo-expander or to further expand the gas mixture to assist in rapid cooling of the membrane device and/or system to the desired operating temperature. In such embodiments, the lower initial productivity of the membrane device may be adequate to deliver small quantities of a product gas. This lower initial productivity may be compensated for by adding a product buffer tank.

In some embodiments, a recirculation of part or all of the permeate gas or the non- permeate gas is desired. For example, an intermediate stream from one membrane device may be recompressed and used as part of the feed stream to the same or another membrane device. In a preferred embodiment, an intermediate stream may be recycled from a set of membrane devices in series or in cascade. In such embodiments, higher recovery of a desired product or higher purity could be achieved.

In other embodiments, a turbo-expander is used for cooling or as a means for recovering a portion of the pressure energy of a product gas. The terms turbo-expander and turbine expander are used interchangeably in this specification. A turbo-expander is a device in which gas is expanded by a turbine, extracting energy in the form of work from the gas stream while reducing the pressure of the gas stream. When energy is extracted from the gas stream in this manner, the gas stream cools significantly, thereby providing particularly effective cooling for the process. The turbo-expander may be used to operate a fan, compressor, pump, generator or other device through a mechanical, hydraulic or other connection. However, in many small plants it may not be economical to recover the energy and the turbo-expander can only drive a braking device. Even in this case, turbo-expanders are a particularly effective cooling device.

In some embodiments, minimizing energy requirements is achieved through the use of an integrated Process design desired. Energy can be saved by using one or more cold streams to cool a warmer stream in a heat exchanger. In one embodiment, cold product streams exiting the membrane device are used to partially cool the feed gas mixture and an external cooling supply is used to further cool the feed gas mixture. In another embodiment, product from a membrane device is cooled by passing it through an expansion device and the cold product gases are used to cool the feed gas mixture in a heat exchanger.

Optionally, the process and apparatus may employ means for changing the temperature and/or pressure of a portion or all of the permeate and/or non-permeate streams exiting the membrane device to a temperature and/or pressure at which the streams are useful. For example, the permeate or non-permeate stream may be too cold to be used effectively, or it may be desirable to recover the cooling potential of the cold stream by using it to cool a warm stream. Or it may be desirable to cool a portion or all of the permeate or non-permeate stream before recycling the stream to an earlier point in the process. Heat exchangers, compressors, and the like may be used for these purposes.

In some embodiments of the invention, it may be desirable to use membranes in combination with other separation processes, such as pressure swing absorption, temperature swing absorption, fractionation, distillation, rectification, chemical reaction, absorption or extraction. For example, membranes may be used to enrich either oxygen or nitrogen in a gas mixture prior to feeding the gas to or from a cryogenic distillation process. Alternatively, a reaction or other process may be used to remove residual impurities from an enriched gas produced via a membrane process. Figure 11 shows an apparatus for separating a gas mixture. A feed gas mixture is introduced through a line 1102 into a first compressor 1104. In the illustrated embodiment, four compressor stages are used. Additional compressor stages or fewer compressor stages may be used as needed to obtain the desired pressure for the feed gas mixture in the separation processes. The gas in line 1102 should preferably be pretreated to remove particulate matter and unwanted impurities using a coalescer or other filter media. The first compressor 1104 is driven by a motor 1106 via a shaft 1107. A single motor may be used to drive a series of compressors represented by a second compressor 1120, a third compressor 1124, and a fourth compressor 1142, each driven by shafts 1110, 1122, and 1125. Gas exits the first compressor 1104 via line 1108 into a cooler 1112 and then exits via line 1113 into a water separator or drift eliminator. Condensed water exits the drift eliminator 1114 via line 1116, and gas exits via line 1118.

The gas from line 1118 enters the second compressor 1120 and exits via line 1130 into a cooling device 1132. The gas exits the cooling device 1132 via line 1134 and then enters the dropper eliminator 1136. Water exits the eliminator 1136 via line 1139 and relatively dry gas exits via line 1138 into a third compressor 1124. The gas exiting the third compressor 1124 flows through line 1126 into a cooler 1140. The gas is then conveyed through line 1149 into the eliminator 1146 with liquid water flowing through line 1148 and relatively dry gas flowing through line 1141. The gas from line 1141 enters the fourth compressor 1142 and exits via line 1143 which conveys gas into the cooler 1144. Gas exits the cooler 1144 via line 1145 and enters the mist eliminator 1150. Liquid water exits the mist eliminator 1150 via line 1147. Compressed gas exits the mist eliminator 1150 via line 1151. Additional drying stages may be used if required, depending on the amount of water in the feed gas mixture and the desired dew point.

From line 1151, the feed gas mixture flows into a cooled cooler 1152 and then via line 1153 into a droplet separator 1154. (Alternatively, the feed gas mixture may be cooled using an external cooling system or by heat exchange with some or all of the permeate gas.) Before entering the cooler 1152, the feed gas mixture is preferably filtered to remove oil picked up by the compressor. Liquid water exits via line 1156 and nearly dry gas flows through line 1158. To achieve additional drying, either an air dryer operating with a molecular sieve or a recuperative heat exchanger may be used. A molecular sieve is shown in Fig. 11 with two sieve beds 1168 and 1170 are shown which are alternately regenerated and operated. The sieve beds may be regenerated using part or all of the permeate gas as a dry sweep, and/or heat regeneration of the sieve beds may also be used. The use of a recuperative heat exchanger would require some modifications to the apparatus shown, but the modifications required would be obvious to one skilled in the art.

Switching valves 1176 and 1160 are used to switch between the various gas flows required for drying and regeneration. Gas from line 1158 flows into a switching valve 1160 and is then directed to one of the molecular sieve beds 1168 and 1170, respectively. For purposes of explanation, the apparatus will be described with bed 1168 being the active drying bed while 1170 is being regenerated. In this case, the dry feed gas mixture flows through valve 1160 into line 1162 and then into the molecular sieve bed 1168. The gas exits through tube 1172 into switching valve 1176 where the gas is directed into tube 1178.

The sieve bed 1170 is regenerated by means of a gas which is introduced through a line 1180 and is dry relative to the bed. This dry gas from the line 1180 flows through the switching valve 1176 into the line 1174 and is passed through the sieve bed 1170. The moisture purge gas then exits through the line 1166 into the switching valve 1160 where the gas is removed through the line 1164.

The feed gas mixture flows through line 1178 into a heat exchanger 1188, which serves to to further cool the reasonably dry gas by contacting it countercurrently with colder gases in lines 1200 and 1210. The heat exchanger is suitably of conventional design with multiple passes connecting each associated inlet and outlet to maximize heat transfer. The cooled feed gas mixture flows into line 1190 where it is conveyed to a second cooler 1192 where heat exchange occurs with gas from line 1208. Alternatively, the heat exchange capability may be provided by a single heat exchanger instead of both heat exchangers 1188 and 1192. The feed gas mixture is then introduced from line 1194 into a membrane unit 1196. The membrane unit 1196 may comprise one or more membrane elements. Some of the heat exchange occurs in the membrane unit 1196 as the permeate gas is at low pressure and cools by expansion and cools the non-permeate gas stream exiting the membrane unit 1196 through line 1198. The non-permeate reject gas stream in line 1198 is still at a relatively high pressure and the gas is cooled by passing into a turbo expander 1206. The energy from the gas expansion is conveyed by shaft 1204 to an energy consuming device 1202 which allows the use of this energy elsewhere or simply converts the energy to heat.

The expanded non-permeate gas is conveyed through line 1208 to a heat exchanger 1192 where it is used to cool the feed gas mixture introduced through line 1190. The gas permeating through membrane 1196 is conveyed through line 1200 to a heat exchanger 1188 where this gas, together with the gas exiting through line 1210 from heat exchanger 1192 is used to cool the feed gas mixture in the heat exchanger 1188. All lines are preferably insulated to preserve the cooling value of the gas contained in the lines where such gases are below ambient temperature.

The gas from line 1200 exits heat exchanger 1188 through line 1214, and a portion of this gas is diverted to line 1184 through valve 1186. The gas that is not diverted to line 1184 is diverted to line 1216. The gas through line 1184 enters an aerator, fan, or other low pressure device for delivering gas 1182 and is then conveyed to line 1180 where this gas is used to vent moisture from the molecular sieve bed being regenerated, either 1168 or 1170. The gas in line 1180 may optionally be heated to assist in regenerating the molecular sieve bed. The gas from line 1210 exits heat exchanger 1188 through line 1212.

In Fig. 11, the size and preferred design of the compression stages, drying stages, heat exchangers and membrane units depend on the exact composition of the feed gas mixture, the type of membrane used, the desired operating temperature and other separation conditions. Aluminum or aluminum alloys are preferred construction materials for many of the system components that operate at low temperatures. It is left to the expert in the field to determine the optimal conditions depending on the composition of the gas mixture and the desired separation.

Figure 14 shows a simplified schematic representation of another apparatus that can be used to carry out the process of the present invention. A gas mixture to be separated is conveyed in a line 1402 to a compressor 1404, which compresses the gas. The compressed gas exits the compressor 1404 through line 1406 and is conveyed to a heat exchanger 1408. The compressed gas exits the heat exchanger 1408 via line 1410 and is introduced into a membrane unit 1412. The non-permeate gas exits the membrane unit 1412 through line 1422 and the permeate gas exits through line 1414. Conduit 1422 carries the non-permeate gas into a turbine expander 1424 and the expanded and cooled gas exits through conduit 1426. The cooled gas flows from conduit 1426 into heat exchanger 1408 and then exits through conduit 1428. The turbine expander is mechanically or otherwise connected to a compressor 1416 through a power transfer device 1420. The power transfer device 1420 may be a drive shaft. The permeate gas in conduit 1414 enters compressor 1416 and exits as a compressed gas through conduit 1418. The device shown in Figure 14 is an effective means for producing a relatively high pressure product from the permeate gas in a membrane system operating at low temperatures. The output from the turbine expander can also be used to create a vacuum on the permeate side of the membrane device, to compress the feed gas mixture or non-permeate gas to a higher pressure, or to do the like.

Fig. 15 shows a simplified schematic view of another device for separating a gas mixture to produce a permeate gas of relatively high purity and re relatively high pressure. The feed gas mixture is introduced into a first compression device 1504 via a line 1502 and then exits through a line 1506. A line 1506 carries the compressed gas to a second compression device or second stage compression device 1508. The compressed gas exits the compression stage 1508 via a line 1510 and is conveyed to a heat exchanger 1512. The cooled gas exits the heat exchanger 1512 via a line 1514 and is conveyed to a first membrane device 1516. The non-permeate gas from the first membrane device 1516 exits through a line 1532 and is conveyed to a turbine expansion device 1534. The turbine expander 1534 expands and cools the gas, and the gas exits through a line 1536, is conveyed to the heat exchanger 1512, and finally exits through a line 1538. If high pressure flow of the non-permeate gas is not desired, the gas can be further expanded so that significant cooling is achieved and significant work is provided. The permeate gas from the first membrane device 1516 exits through a line 1518 and is conveyed to a second membrane device 1520. Gas which does not permeate the second membrane device 1520 exits through a line 1530 and is conveyed to the heat exchanger 1512. The non-permeate gas from the second membrane device 1520 is conveyed from the heat exchanger 1512 through a line 1540 into a line 1506 and then into the second compression stage 1508. This allows additional recovery of the desired permeate gas by recirculation. In the second membrane device 1520, the permeate gas exits through a line 1522 and is conveyed to a compression device 1524, which is connected to a turbo-expansion device 1534 by an energy transfer device 1528. The turbo-expansion device 1534 can thereby supplying part or all of the power required for the compression device 1524. The permeate gas from the second membrane device 1520 is conveyed via a line 1526 after compression in the compression device 1524. The resulting product gas in the line 1526 has a relatively high pressure and a relatively high purity.

Figure 16 shows in schematic form another embodiment of the invention which uses conventional refrigeration rather than a turbine expander to achieve the desired temperature for the feed gas mixture to the membrane. In Figure 16, a feed gas mixture is introduced through a line 1602 into a compression device 1604. Gas exits the compression device 1604 via a line 1610. A line 1610 carries gas into a first heat exchanger 1618 and after cooling the gas is carried via a line 1620 into a second heat exchanger 1622. The feed gas from line 1620 exits the heat exchanger 1622 via a line 1624 and is introduced into a membrane unit 1626. The non-permeate gas exits the membrane device 1626 via a line 1628 and is introduced into the first heat exchanger 1618 and finally exits through line 1632. The gas permeating the membrane device 1626 is conveyed via a line 1630 into the first heat exchanger 1618 and exits through a line 1634. The second heat exchanger 1622 is cooled by a heat transfer fluid which is introduced into the heat exchanger 1622 via a line 1638. The heat transfer fluid exits the second heat exchanger 1622 via a line 1640 and is then introduced into a cooling device 1642 in which the heat transfer fluid is cooled and then circulated via a line 1638. Heat transfer capability may alternatively be provided by a single Heat exchangers can be supplied instead of the two heat exchangers 1618 and 1622 shown.

One skilled in the art will recognize that a variety of alternatives can be used for changing the temperature of the gas mixture to operating temperature, removing contaminants, constructing membranes, and compressing or expanding gases. For example, a recuperative heat exchanger can be used to remove certain condensable contaminants from a feed gas mixture. The term membrane unit or membrane device as used herein refers to multiple membranes arranged in series, parallel, cascade or a recycle arrangement so as to increase the product purity or productivity of the system. The various devices described herein in Figures 11, 14, 15 and 16 are shown for illustration purposes and are not intended to unnecessarily limit the devices used in the various stages of the process which is the subject of the invention.

The membrane, the gas mixture to be separated, the separation temperature and the transmembrane pressure should be selected to achieve an economically efficient separation of component gases in the feed gas mixture. It is desirable that the permeability of the component gas which selectively permeates through the membrane be sufficiently high under the selected conditions so that the membrane device does not require excessive membrane surface area. Preferably, for a separation of oxygen and nitrogen, the permeability of oxygen under the selected conditions is at least 0.1 Barrers, more preferably at least 1 Barrer. The permeability of oxygen is preferably less than 2000 Barrers at 30°C since transport through high permeability membranes generally exhibits anomalies. which cannot be directly attributed to a solution-diffusion mechanism, and therefore such membranes are not relevant for the process which is the subject of the invention.

Figure 13 shows the performance of membranes made from certain polymers that meet the separation properties of Equation I, as well as the performance of other polymers that do not meet these criteria. The membranes according to points 1 to 10 consist of the following polymers: Point 1 is poly- 4-methylpentene-1 (PMP), point 2 is a polymer which is derived from 4, 6 - diaminoresorcinol dihydrochloride and 1,1,3-trimethyl-3-phenyl-indane-4,5'-dicarboxylic acid (PBO A), point 3 is polyphenylene oxide (PPO), point 4 is tetrachloro- hexafluoro bisphenol A polycarbonate (TCHF BA PC), point 5 is a polymer which is derived from 2,2-bis(3-amino-4- hydroxyphenyl)-1,1,1,3,3,3-hexafluoropropane and diphenyl ether- 4,4'-dicarboxylic acid (PBO B), point 6 is tetrabromohexafluoro bisphenol A isophthalate ester (TBHF BA IE), point 7 is Cellulose Triacetate (CTA), Item 8 is TRYCITE Polystyrene (PST), Item 9 is Tetrabromo Bisphenol A Polycarbonate (TBBA PC), and Item 10 is 9,9-bis(3,5-dibromo-4-hydroxy phenyl)fluorene Polycarbonate.

SPECIFIC IMPLEMENTATION EXAMPLES

The subject matter of the invention is further illustrated by the following examples, but is not limited thereby.

Process simulation examples 1-10.

Ten examples describing low temperature membrane processes are presented to demonstrate the scope of the invention. These examples summarize process simulations of based on commercially available process engineering computer software and known membrane device parameter relationships, using hollow fiber membrane devices and assuming ideal countercurrent fluid flow patterns and adiabatic behavior with ideal heat transfer between the permeate and non-permeate gas streams. Mass transfer across the membrane was described using standard gas separation equations in terms of partial pressure driven permeability, with temperature dependent permeabilities derived from data in Tables 13 and 15. Recovery, as used herein, refers to the fraction or percentage of the volume of product gas (permeate or non-permeate gas, respectively) to the volume of feed gas. The examples include a process flow diagram, a main stream summary, and a description of the main process equipment. Each example describes a process which is considered to be operable.

The examples are as follows:

1. Nitrogen process with an expansion device on the non-permeate product stream.

2. Nitrogen process with an expansion device on the permeate product stream.

3. Nitrogen process with an external cooling system.

4. Nitrogen process with an external cooling system and a dehydration membrane for air drying.

5. Oxygen process with an oxygen compression device operated by an expansion device.

6. Oxygen process with a vacuum pump on the permeate product stream.

7. Oxygen process with cascaded membrane devices and a recycle stream.

8. Methane purification process using an expansion device on the non-permeate product stream.

9. Methane purification process with an external cooling system.

10. Carbon dioxide recovery process without expansion device or external cooling.

Examples 1-7 are for the separation of oxygen and nitrogen from air using PBO A membranes, and Examples 8-10 are for the separation of methane and carbon dioxide mixtures using tetrachlorohexafluoro bisphenol A polycarbonate (TCHF BA PC) membranes. Examples 1-3 demonstrate the use of a low temperature membrane process to produce a nitrogen-enriched non-permeate product using three different methods for removing heat from the process to maintain a low temperature. Example 4 is similar to Example 3, except that a membrane device which is selectively permeable to water is used to add nitrogen to the feed. gas air stream to remove water vapor, thereby preventing condensation in the process. This example also demonstrates the use of membranes operating at ambient temperature in conjunction with low temperature membranes. Examples 5-7 demonstrate the use of low temperature membrane processes to produce an oxygen-enriched permeate product. Example 5 demonstrates the use of a compression device to maintain the pressure differential across the membrane device and the use of an expansion device to operate a product compression device. Example 6 demonstrates the use of a vacuum pump in conjunction with a blower to maintain a pressure differential across the membrane device. Example 7 is similar to Example 5 except that cascaded membrane devices and a recycle stream are used to achieve higher product purity. Examples 8-10 demonstrate the use of a low temperature membrane carbon dioxide/methane separation process when a contaminant-free feed gas stream is available at high pressures and ambient temperatures. Example 8 uses heat exchangers and an expander to change the temperature of the streams and maintain the membrane device at a lower temperature where a higher pressure non-permeate product stream is desired. Example 9 is similar to Example 8 except that the expander is replaced by an external cooling system. Example 10 shows how the self-cooling capacity of a gas can be used for cooling, eliminating the need for an expander or external cooling system to remove heat from the process.

EXAMPLE 1 NITROGEN PROCESS WITH AN EXPANSION DEVICE IN THE NON-PERMATE PRODUCT STREAM

Example 1 is a process for producing an enriched nitrogen product stream from air. The nitrogen stream specifications are a flow rate of 40,000 standard cubic feet per hour (SCFH) (1,130 cubic meters per hour) at 125 psig (863 kPa) pressure with a maximum 1 percent oxygen concentration on a molar basis. A simplified flow diagram for this process is shown in Figure 1. A summary of the major process streams is given in Table 1A and a summary description list of the major process equipment is given in Table 1B. Definitions of abbreviations used are given in Table 1C.

In Fig. 1, process feed air 110 is available from the atmosphere at ambient pressure, temperature and humidity and must be compressed. A four stage centrifugal compressor 105 with intercoolers and aftercoolers is shown. Some water condensation may occur in the coolers, so the coolers must be provided with automatic drains 120 AD to remove the condensed liquid. Air 110 entering compressor 105 is normally filtered to remove particles which could erode the compressor elements. Air 112 leaving compressor 105 is passed through a small heat exchanger 121 where it is cooled to between 4ºC and 10ºC with evaporative coolant 122 from an external cooling source. This causes an additional condensation of water from the air, which is separated in a container 123, which is provided with an automatic drain 124. If the compressor is an oiling type, entrained or evaporated oil is also of the air in a vessel 123. Further drying of the air 113 takes place in an adsorption air dryer 126. In this example, a double layer molecular sieve type 3A adsorber available from Union Carbide is used, with one layer being used while the other layer is regenerated. This adsorber 126 dries the incoming air 113 to a dew point below the lowest temperature expected to be reached in the process, typically a dew point of less than -100°C. Not shown in the flow diagram is additional purification equipment used to remove further contaminants from the air stream. Such equipment may include converging filters to remove runoff liquid droplets after the water separation tank 123, an activated carbon filter to remove hydrocarbons which could have a damaging impact on the membrane and/or fine particle filters to remove carbon powder or dust from the air after the adsorber 126.

The compressed and cleaned air 130 enters a main heat exchanger 127 where it is cooled with heating product streams 153 and 140. The air 131 is further cooled in a small heat exchanger 128 with the heating product gas 152 exiting the expansion device 129. The cooled air 132 then enters the high pressure passages of the membrane device 133 and the oxygen in the air preferentially permeates through the membrane 134 to the low pressure passages of the membrane device. This results in the high pressure non-permeate stream 150 exiting the membrane device 133 enriched in nitrogen while the low pressure permeate stream 140 exits the membrane device 133 enriched in oxygen. The non-permeate product stream 151 enters the expansion device 129, which removes some energy from the jet, thereby increasing its pressure and temperature. temperature. The exit stream from the expander 129 is 152. This removal of energy offsets the heat entering the cold equipment through heat transfer with the environment and other sources, thus maintaining the equipment and process streams at the desired low temperature. The product streams 152 and 153 then enter heat exchangers 128 and 127, respectively, where the streams are heated with cooling feed air streams 131 and 130, respectively. Thereafter, the high pressure non-permeate nitrogen product stream 154 enters a pipeline while most of the low pressure permeate oxygen product stream 141 is simply vented to the atmosphere through a line 142. A small amount of the dry low pressure oxygen product 143 may be compressed and heated and then used to regenerate the inactive layer of the adsorption air dryer 126 via lines 143 and 144.

This example demonstrates many of the features of a low temperature membrane process. A number of membrane modules are arranged in parallel to achieve the desired separation at the given pressures. A compressor is used to maintain the pressure difference across the membrane. The cold process equipment and piping are well insulated to minimize heat leakage into the system. An expansion device is used to provide cooling and to maintain the cold process equipment at the desired temperature. A refrigerated air dryer, an adsorption air dryer, an activated carbon filter, a converging filter and particulate filters are used to remove contaminants from the air. Heat exchangers are used to cool the feed gas stream and heat the product gas streams. By exchanging heat between these streams, the required cooling can be provided by the expansion ation device, thereby saving energy. No product recompression is necessary in this process because the air compressor is designed to compensate for the pressure drop in the system so that the nitrogen product leaving the main heat exchanger is at the specified pressure. Table 1A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM Table 1A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM Table 1A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM Table 1A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM

TABLE 1B LOW TEMPERATURE NITROGEN PROCESS WITH EXPANSION DEVICE ON A NON-PERMATE STREAM: MAIN EQUIPMENT SUMMARY MAIN AIR COMPRESSOR

4-stage compressor with intercoolers and aftercooler. Capacity: 0.73 kg/s air at 1284000 Pa discharge pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 242 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

MEMBRANE MODULES

13.8 parallel modules containing 1.4 x 106 PBO-A fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 1.5748

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Three groups of passages in counterflow arrangement.

Power, current 1: 0.71 kg/s at 1200700 Pa from 4.3ºC to -56ºC.

Stream 2: 0.37 kg/s at 982500 Pa from -63ºC to - 0.4ºC.

Stream 3: 0.34 kg/s at 120700 Pa from -60ºC to - 0.4ºC.

Useful power: 44.4 kW, maximum 4.7ºC warm end temperature difference.

Maximum pressure drop: 20700 Pa at any flow.

SECONDARY COOLER HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Two sets of passages in counterflow arrangement.

Power, current 1: 0.71 kg/s at 1180000 Pa from -56ºC to -60ºC.

Stream 2: 0.37 kg/s at 1032000 Pa from -70ºC to -63ºC.

Useful power: 2.9 kW, maximum 3.0ºC warm end temperature difference.

Maximum pressure drop: 20700 Pa at any flow.

EXPANSION DEVICE

Brake-loaded turbo expansion device for extracting 2.35 kW of power.

Performance: 0.37 kg/s, 99 percent nitrogen at 1143000 Pa, -64ºC.

Pressure reduction to 1003000 Pa.

Minimum efficiency: 80 percent isentropic.

COOLER/DRYER

Heat exchanger with external Freon cooling system and water separation tank.

Performance: 0.73 kg/s at 1283000 Pa from 40ºC to 4.4ºC.

Expected cooling capacity: 36 kW.

Maximum total pressure drop: 13800 Pa.

AIR DRYER

Double vessel molecular sieve adsorption air dryer with regenerative blower and heater.

Capacity: 0.73 kg/s saturated air at 1270000 Pa, 4.4ºC.

Dew point requirement: -70ºC at pressure.

Maximum pressure drop: 69000 Pa

Table 1C: Equipment list explanation for Table 1A

COMPRESS Main air compressor

AFT-COOL compressor aftercooler

CHILLER cooled air cooling device

DRYER molecular sieve adsorption air dryer

HTEX main heat exchanger

SUBC small heat exchanger

MEMBRANE Parallel membrane modules

HLEAK heat leak simulation

EXPAND expansion device

EXAMPLE 2 NITROGEN PROCESS WITH AN EXPANSION DEVICE IN THE PERMATE PRODUCT STREAM

Example 2 is also a process for producing an enriched nitrogen product stream from air. The nitrogen jet specifications are the same as Example 1. A simplified flow diagram for this process is shown in Figure 2. A summary of the major process streams is given in Table 2A and a summary description list of the major process equipment is given in Table 2B. Definitions of abbreviations used are given in Table 2C.

This process is similar to that described in Example 1, however, the expansion device 229 is located on the low pressure oxygen-enriched permeate product stream 240 exiting the membrane device 233 rather than on the high pressure nitrogen-enriched non-permeate product stream 250. The small heat exchanger between the main heat exchanger 227 and the membrane device 233 has also been eliminated. Due to these changes, the pressure differential across the membrane 234 becomes smaller, thus more parallel membrane modules are required. A small amount of the dry low pressure oxygen product 243 can be compressed and heated and then used to regenerate the inactive layer of an adsorption dryer 226 via lines 243 and 244. The expansion device 229 would need to be larger to accommodate a lower pressure product stream. The size of other process equipment can also vary to accommodate different stream conditions. Consequently, the capital costs for this process are higher than those for Example 1. The recovery of the process is lower, so more air must be fed into the membrane device 233 to produce the same amount of nitrogen. However, the total pressure of the air leaving the compressor 205 is lower, so the energy cost for this process is slightly less than for Example 1.

The process demonstrates the same features as Example 1. It also demonstrates flexibility in selecting the location of process equipment and the arrangement of process streams. This example further demonstrates how rearranging a piece of process equipment affects the size and power requirements of other equipment and therefore the economics of the process. Table 2A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM Table 2A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM Table 2A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM Table 2A STREAM SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXPANSION DEVICE ON THE PERMATE PRODUCT STREAM

TABLE 2B LOW TEMPERATURE EXPANSION NITROGEN PROCESS ON PERMATE STREAM: MAIN EQUIPMENT SUMMARY MAIN AIR COMPRESSOR

4-stage compressor with intercoolers and aftercooler.

Capacity: 0.76 kg/s air at 1100600 Pa delivery pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 237 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

MEMBRANE MODULES

18.5 parallel modules containing 1.4 x 106 PBO-A fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 1.5748

Tube sheet length (m): 0.2032

HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Three groups of passages in counterflow arrangement.

Power, current 1: 0.74 kg/s at 1087000 Pa from 4.4ºC to -60ºC.

Stream 2: 0.37 kg/s at 982500 Pa from -64ºC to 1.4ºC.

Stream 3: 0.37 kg/s at 120700 Pa from -64ºC to 1.4ºC.

Useful power: 49.1 kW, maximum 3.0ºC warm end temperature difference.

Maximum pressure drop: 20700 Pa at any flow.

EXPANSION DEVICE

Brake-loaded turbo expansion device for extracting 1.53 kW of power.

Performance: 0.37 kg/s, 42 percent oxygen at 131600 Pa, -60ºC.

Pressure reduction to 120700 Pa.

Minimum efficiency: 80 percent isentropic.

COOLER/DRYER

Heat exchanger with external Freon cooling system and water separation tank.

Performance: 0.76 kg/s at 110060 Pa from 40ºC to 4.4ºC.

Expected cooling capacity: 35 kW.

Maximum total pressure drop: 13800 Pa.

AIR DRYER

Double vessel molecular sieve adsorption air dryer with regenerative blower and heater.

Capacity: 0.75 kg/s saturated air at 109000 Pa, 4.4ºC.

Dew point requirement: -70ºC at pressure.

Maximum pressure drop: 69000 Pa

Table 2C: Equipment list explanation for Table 2A

COMPRESS Main air compressor

AFT-COOL compressor aftercooler

CHILLER cooled air cooling device

DRYER molecular sieve adsorption air dryer

HTEX main heat exchanger

SUBC small heat exchanger

MEMBRANE Parallel membrane modules

HLEAK heat leak simulation

EXPAND expansion device

EXAMPLE 3 NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM

Example 3 is also a process for producing an enriched nitrogen product stream from air. The nitrogen stream specifications are the same as for Example 1. A simplified flow diagram for this process is shown in Figure 3. A summary of the major process streams is given in Table 3A and a summary description list of the major process equipment is given in Table 3B. Definitions of abbreviations used are given in Table 3C.

This process is similar to that described in Example 1. A feed gas 310 passes through the components of a condenser 305 as stream 311. However, the expansion device is removed and the small heat exchanger between the main heat exchanger 327 and the membrane device 333 is replaced by an evaporative condenser 364 in which the incoming air 331 in one set of passes is cooled by an incoming boiling refrigerant in a line 362 in the other set of passes. A cold liquid refrigerant is supplied by an external cooling system 365 which is here represented by a simple single stage arrangement. In this arrangement the refrigerant in a line 363 is compressed by a compressor 366 to a high pressure in a line 360, then condensed in a heat exchanger 367 and subjected to exchange with ambient air. The condensed refrigerant in a line 361 is expanded to a low pressure by a throttle valve 368. The partially evaporated refrigerant is passed through a line 362 into an evaporator 364 where an exchange thereof with a feed gas in a line 331. The evaporative coolant in a line 363 exiting the evaporator 364 is returned to the coolant compressor 366.

Because of these changes, the pressure differential across the membrane 334 is smaller than in Example 1, so more parallel membrane modules are required, but not as many as in Example 2. Due to changes in process recovery and air compressor pressure, the cost of the main air compressor 305 for this process is lower than for Examples 1 and 2. However, there is an additional energy cost for the cooling compressor 366.

This process has most of the same features as Example 1. It also illustrates how the expansion device can be replaced by an external cooling system to maintain the process equipment at the desired temperature. External cooling systems are well known and many options are available. For example, the simple single-stage cooling system illustrated can be replaced by a multi-stage system. Such modifications to external cooling systems are within the scope of Example 3.

Table 3D summarizes the simulation data generated by Examples 1-3, which are described in Tables 1A through 3C. Published literature performance data for pressure swing adsorption (PSA) typically report energy/100 SCF of 99 percent nitrogen product (120 psig (827 kPa) delivery pressure) between 0.75 and 1.2 kWh/100 SCF of product, depending on system size. The data presented in Table 3D indicate a clear performance improvement over current PSA technology in terms of energy requirements. Table 3A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM Table 3A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM Table 3A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM Table 3A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM

TABLE 3B LOW TEMPERATURE NITROGEN PROCESS WITH EXTERNAL COOLING: MAIN EQUIPMENT SUMMARY MAIN AIR COMPRESSOR

4-stage compressor with intercoolers and aftercooler.

Capacity: 0.75 kg/s air at 1115000 Pa delivery pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 236 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

MEMBRANE MODULES

17.4 parallel modules containing 1.4 x 106 PBO-A fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 1.5748

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Three groups of passages in counterflow arrangement.

Power, current 1 : 0.73 kg/s at 1032000 Pa from 4.4ºC to -58ºC.

Stream 2: 0.37 kg/s at 982500 Pa from -64ºC to 1.4ºC.

Stream 3: 0.36 kg/s at 120700 Pa from -60ºC to 1.4ºC.

Useful power: 46.7 kW, maximum 3.0ºC warm end temperature difference.

Maximum pressure drop: 20700 Pa at any flow.

COOLER/DRYER

Heat exchanger with external Freon cooling system and water separation tank.

Performance: 0.74 kg/s at 1115000 Pa from 40ºC to 4.4ºC.

Expected cooling capacity: 34 kW.

Maximum total pressure drop: 13800 Pa.

AIR DRYER

Double vessel molecular sieve adsorption air dryer with regenerative blower and heater.

Capacity: 0.74 kg/s saturated air at 1101000 Pa, 4.4ºC.

Dew point requirement: -70ºC at pressure.

Maximum pressure drop: 69000 Pa

EXTERNAL COOLING COMPRESSOR

Freon-13 reciprocating compressor with a capacity to compress 0.058 kg/s Freon-13 vapor at 3745000 Pa, 97ºC to liquid at 27ºC with ambient air at 20ºC.

Expected net power: 6.28 kW.

EXTERNAL COOLING CONDENSER

Heat exchanger with a capacity to condense 0.058 kg/s Freon-13 vapor at 3745000 Pa, 97ºC to liquid at 27ºC with ambient air at 20ºC.

Expected net power: 6.28 kW

EXTERNAL COOLING EVAPORATOR

Evaporative heat exchanger with a capacity to cool 0.73 kg/s of air at 1012000 Pa from -58ºC to -60ºC with Freon-13 boiling at -66ºC.

Expected net power: 1.5 kW.

Table 3C: Equipment list explanation for Table 3A

COMPRESS Main air compressor

AFT-COOL compressor aftercooler

CHILLER cooled air cooling device

SEPARAT Molecular Sieve Adsorption Air Dryer

HTEX main heat exchanger

EVAP External Cooling Evaporator

MEMBRANE Parallel membrane modules

HLEAK heat leak simulation

RCOMPR External cooling compressor

RCNDNS External cooling condenser

RJT External Cooling Expansion Valve TABLE 3D SUMMARY OF NITROGEN PROCESSES

EXAMPLE 4 NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM AND DEHYDRATION MEMBRANES FOR AIR DRYING

Example 4 is also a process for producing an enriched nitrogen product stream from air. The specifications of the nitrogen stream are the same as for Example 1. A simplified flow diagram for this process is shown in Figure 4. A summary of the major process streams is given in Table 4A and a summary description list of the major process equipment is given in Table 4B. Definitions of abbreviations used are given in Table 4C.

This process is similar to that described in Example 3. However, the refrigerated air cooler and separation vessel between the air compressor 405 and the adsorption air dryer 426 have been replaced by a set of parallel membrane devices 414 which preferentially pass water while essentially not passing other air constituents. A larger portion of the dry oxygen-enriched permeate stream 472 must be slightly compressed to act as a sweep gas in the low pressure passages of the dehydration membrane devices 414. Due to the fact that this device is expected to dry the air to a dew point well below 0°C, the adsorption air dryer 426 would be considerably smaller. However, due to the fact that the air is no longer cooled at this point, the main heat exchanger 427 must be somewhat larger. The remainder of the process equipment in this Example is similar to the process equipment in Example 3.

This example demonstrates the variety of choices available for removing contaminants from the process feed gas. It also illustrates how membrane devices can be operated in series and at different temperatures in a low temperature membrane process. Table 4A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM AND DEHYDRATION MEMBRANES FOR AIR DRYING Table 4A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM AND DEHYDRATION MEMBRANES FOR AIR DRYING Table 4A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM AND DEHYDRATION MEMBRANES FOR AIR DRYING Table 4A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM AND DEHYDRATION MEMBRANES FOR AIR DRYING Table 4A POWER SUMMARY FOR A LOW TEMPERATURE NITROGEN PROCESS WITH AN EXTERNAL COOLING SYSTEM AND DEHYDRATION MEMBRANES FOR AIR DRYING

TABLE 4B LOW TEMPERATURE NITROGEN PROCESS WITH EXTERNAL COOLING: DEHYDRATION MEMBRANE

Air Dryer: Main Equipment Summary

MAIN AIR COMPRESSOR

4-stage compressor with intercoolers and aftercooler.

Capacity: 0.74 kg/s air at 1070000 Pa delivery pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 228 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

MEMBRANE MODULES

17.4 parallel modules containing 1.4 x 106 PBO-A fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 1.5748

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Three groups of passages in counterflow arrangement.

Power, current 1: 0.73 kg/s at 1032000 Pa from 40ºC to -58ºC.

Stream 2: 0.37 kg/s at 982500 Pa from -63ºC to 37ºC.

Stream 3: 0.36 kg/s at 120700 Pa from -60ºC to 37ºC.

Useful power: 73.1 kW, maximum 3.0ºC warm end temperature difference.

Maximum pressure drop: 20700 Pa at any flow.

DEHYDRATION MEMBRANES

10 spiral wound modules containing 20 perfluorosulfonic acid (PFSA) membrane sheets.

Blade length (m): 0.5

Sheet thickness (m): 5.08 x 10&supmin;&sup5;

Blade width (m): 1.0

DEHYDRATION DIAPHRAGM BLOWER

Centrifugal fan with aftercooler. Engine power: 3.6 kW Braking power.

Output: 0.1 kg/s 43% O2 at 131000 Pa discharge pressure, 40ºC.

AIR DRYER

Double vessel molecular sieve adsorption air dryer with regenerative blower and heater.

Capacity: 0.73 kg/s air at 1101000 Pa, 40ºC with a dew point of -17ºC.

Dew point requirement: -70ºC at pressure.

Maximum pressure drop: 27600 Pa

EXTERNAL COOLING COMPRESSOR

Freon-13 reciprocating compressor with a capacity to compress 0.068 kg/s from 202430 Pa to 3745000 Pa discharge pressure.

Engine power: 5.6 kW braking power.

EXTERNAL COOLING CONDENSER

Heat exchanger with a capacity to condense 0.068 kg/s of Freon-13 vapor at 3745000 Pa, 97ºC to a liquid at 27ºC with an ambient air of 20ºC.

Expected net power: 7.4 kW.

EXTERNAL COOLING EVAPORATOR

Evaporative heat exchanger with a capacity to cool 0.73 kg/s of air at 1012000 Pa from -58ºC to -60ºC with Freon-13 boiling at -66ºC.

Expected net power: 1.8 kW.

Table 4C: Equipment list explanation for Table 4A

COMPRESS Main air compressor

AFT-COOL compressor aftercooler

CHILLER cooled air cooling device

DEHYDRATE dehydration membrane modules

DRYER molecular sieve adsorption air dryer

HTEX main heat exchanger

EVAP External Cooling Evaporator

MEMBRANE Parallel membrane modules

HLEAK heat leak simulation

RCOMPR External cooling compressor

RCNDNS External cooling condenser

RJT External Cooling Expansion Valve

SPLITTER exhaust flow divider

BLOWER Dehydration purge gas blower

BCOOLER fan aftercooler

EXAMPLE 5 OXYGEN PROCESS WITH A PRODUCT COMPRESSOR DRIVEN BY AN EXPANSION DEVICE

Example 5 is a process for producing an enriched oxygen product stream from air. The oxygen stream specifications are a flow rate of 10,000 standard cubic feet per hour (SCFH) (283 cubic meters per hour) at a pressure of 40 psia (276 kPa) with an oxygen concentration of 65 percent on a molar basis. A simplified flow diagram for this process is shown in Figure 5. A summary of the major process streams is given in Table 5A and a summary description list of the major process equipment is given in Table 5B. Definitions of abbreviations used are given in Table 5C.

Process feed air 510 is available from the atmosphere at ambient pressure, temperature and humidity and must be compressed. The air compression and air purification section 505, 521, 523 and 526 of this process is similar to that of Example 1. However, a cold exhaust stream 572 is available from this process and this can be used to replace the coolant in the air cooler 521 between the feed air compressor 505 and the adsorption dryer 526. The compressed, purified air stream 530 is fed into a small heat exchanger 527 where it is cooled with a heating section of the nitrogen-enriched non-permeate exhaust stream 553. The cooled air 531 is then introduced into various parallel membrane devices 533 where separation takes place. The nitrogen-enriched high pressure non-permeate stream leaving the membrane devices 533 via line 550/552, is then expanded to near ambient pressure 569 in the expander 529 and a portion of this stream 553 enters a set of passes in the heat exchanger 527 where it is cooled with the cooling feed air stream 530. The stream 555 is recombined with the colder portion 554 that did not pass through the heat exchanger 527 and most of the stream 570 is vented to the atmosphere. A portion of the dry, cold, nitrogen-enriched non-permeate exhaust stream 571 may be used to regenerate one layer of the adsorption dryer 526 and another portion 572 may be used to compress the air stream 512 after the feed compressor 505. The oxygen-enriched low pressure non-permeate product 540 from the membrane device 533 is fed to a product compressor 574 where it is compressed to the desired product pressure 541. This compressor may be driven by the expander 529 via a shaft or belt drive 575, or the expander may drive an electrical generator 576.

This example demonstrates both the use of a low temperature membrane process to produce an oxygen-enriched permeate product stream and how the permeate product can be recompressed to a higher pressure to be useful. This example further demonstrates that a process does not need to be carefully designed to reduce the cooling requirement. Due to the fact that 75 percent of the feed flow of this process is available to be expanded from high pressure to low pressure, a significant excess of cooling is available. Table 5A POWER SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A COMPRESSOR DRIVEN BY AN EXPANSION DEVICE Table 5A POWER SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A COMPRESSOR DRIVEN BY AN EXPANSION DEVICE Table 5A POWER SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A COMPRESSOR DRIVEN BY AN EXPANSION DEVICE Table 5A POWER SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A COMPRESSOR DRIVEN BY AN EXPANSION DEVICE

TABLE 5B LOW TEMPERATURE OXYGEN PROCESS USING AN EXPANSION DEVICE DRIVEN OXYGEN COMPRESSOR: Main Equipment Summary MAIN AIR COMPRESSOR

4-stage compressor with intercoolers and aftercooler.

Capacity: 0.41 kg/s air at 1238000 Pa discharge pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 131 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

MEMBRANE MODULES

7 parallel modules containing 1.4 · 10⁶ PBO-A fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 1.5748

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Two sets of passages in counterflow arrangement.

Power, current 1: 0.40 kg/s at 1155000 Pa from 4.3ºC to -80ºC.

Stream 2: 0.23 kg/s at 122000 Pa from -157ºC to - 5.7ºC.

Useful power: 35.2 kW, maximum 10ºC warm end temperature difference.

Maximum pressure drop: 13800 Pa at any flow.

EXPANSION DEVICE

Generator-loaded turbo expansion device for extracting 17.6 kW of power.

Performance: 0.30 kg/s, 92 percent nitrogen at 1122000 Pa, -80ºC.

Pressure reduction to 122000 Pa.

Minimum efficiency: 80 percent isentropic.

COOLER/DRYER

Heat exchanger, which is cooled by a nitrogen exhaust stream, and water separation tank.

Power, current 1: 0.40 kg/s at 1283000 Pa from 40ºC to 4.4ºC.

Flow 2: 0.30 kg/s at 114000 Pa from -43ºC to 5ºC.

Net power: 18.6 kW.

Maximum total pressure drop: 13800 Pa.

AIR DRYER

Double vessel molecular sieve adsorption air dryer with regenerative blower and heater.

Capacity: 0.40 kg/s saturated air at 1224000 Pa, 4.4ºC.

Dew point requirement: -80ºC at pressure.

Maximum pressure drop: 69000 Pa

OXYGEN COMPRESSOR

Centrifugal compressor with aftercooler.

Output: 0.10 kg/s 65% oxygen at 276000 Pa delivery pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 7.5 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

Table 5C: Equipment list explanation for Table 5A

COMPRESS Main air compressor

AFT-COOL compressor aftercooler

CHILLER cooled air cooling device

SEPARAT Molecular Sieve Adsorption Air Dryer

HTEX main heat exchanger

MEMBRANE Parallel membrane modules

HLEAK heat leak simulation

EXPAND expansion device

N2SPLIT heat exchanger bypass divider

N2MIX heat exchanger bypass mixer

O2COMPR Product Oxygen Compressor

EXAMPLE 6 OXYGEN PROCESS WITH A PERMATE VACUUM PUMP

Example 6 is a process for producing an enriched oxygen permeate stream from air. The oxygen stream specifications are the same as Example 5 except that the stream pressure is at 20 psia (138 kPa). A simplified flow diagram for this process is shown in Figure 6. A summary of the major process streams is given in Table 6A and a summary description list of the major process equipment is given in Table 6B. Definitions of abbreviations used are given in Table 6C.

The equipment arrangement in this process is similar to that of Example 3, however, the high pressure passages of the membrane device 633 are maintained at a substantially lower pressure and the low pressure passages of the membrane device are maintained under vacuum. Thus, the four stage air compressor of Example 3 is replaced here by a single stage compressor 605. Further, the permeate stream 641 exiting the main heat exchanger 627 is fed to a vacuum pump 680, which is shown as a two stage reciprocating piston type. Due to the lower pressure differential across the membrane 634, a significantly larger number of membrane devices must be used in this example than in Example 5. Furthermore, since operation at lower pressure results in a substantially higher pressure drop for the same mass flow rate, much of the process equipment is significantly larger to keep pressure drops small. This includes the air cleaning equipment, heat exchangers and piping. Vacuum pumps are typically significantly smaller. less effective than compressors, which leads to higher energy costs.

This example demonstrates the use of a vacuum pump in conjunction with a compressor to maintain a pressure differential across the membrane devices. It further demonstrates the use of a vacuum pump to change the pressure of a product stream. It further demonstrates the use of a membrane device at a lower pressure differential than is usually preferred. Table 6A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A VACUUM PUMP ON THE PERMATE PRODUCT STREAM Table 6A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A VACUUM PUMP ON THE PERMATE PRODUCT STREAM Table 6A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A VACUUM PUMP ON THE PERMATE PRODUCT STREAM Table 6A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH A VACUUM PUMP ON THE PERMATE PRODUCT STREAM

TABLE 6B LOW TEMPERATURE OXYGEN PROCESS WITH PERMATE VACUUM PUMP: MAIN EQUIPMENT SUMMARY MAIN AIR COMPRESSOR

1-stage compressor with aftercooler.

Capacity: 0.71 kg/s air at 334400 Pa delivery pressure.

Minimum stage efficiency: 72 percent isentropic.

Engine power: 130 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

MEMBRANE MODULES

14 parallel modules containing 1.4 · 10⁶ PBO-A fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 1.5748

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Three groups of passages in counterflow arrangement.

Power, current 1: 0.69 kg/s at 265500 Pa from 4.3ºC to -56ºC.

Stream 2: 0.59 kg/s at 124100 Pa from -64ºC to 0.3ºC.

Current 3: 0.10 kg/s at 20690 Pa from -60ºC to 0.3ºC.

Useful power: 42.1 kW, maximum 4ºC warm end temperature difference.

Maximum pressure drop: 20680 Pa on stream 1 & 2, 13800 Pa on stream 3.

COOLER/DRYER

Heat exchanger with external Freon cooling system and water separation tank.

Performance: 0.70 kg/s at 334500 Pa from 43ºC to 4.4ºC.

Expected cooling capacity: 40 kW.

Maximum total pressure drop: 13800 Pa.

AIR DRYER

Double vessel molecular sieve adsorption air dryer with regenerative blower and heater.

Capacity: 0.69 kg/s saturated air at 320700 Pa, 4.4ºC.

Dew point requirement: -70ºC at pressure.

Maximum pressure drop: 34500 Pa

VACUUM PUMP

Two-stage reciprocating piston vacuum pump with intercooler and aftercooler.

Output: 0.10 kg/s 65% oxygen at 131000 Pa delivery pressure.

Feed pressure: 6900 Pa.

Engine power: 50.6 kW braking power assuming a polytropic efficiency of 72 percent.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

EXTERNAL COOLING COMPRESSOR

Freon-13 reciprocating compressor with a capacity to compress 0.10 kg/s from 202430 Pa to 3745000 Pa discharge pressure.

Engine power: 8.25 kW braking power.

EXTERNAL COOLING CONDENSER

Heat exchanger with a capacity to condense 0.10 kg/s of Freon-13 vapor at 3745000 Pa, 97ºC to a liquid at 27ºC with an ambient air of 20ºC.

Expected net power: 11.0 kW.

EXTERNAL COOLING EVAPORATOR

Evaporative heat exchanger with a capacity to cool 0.69 kg/s of air at 265500 Pa from -58ºC to -60ºC with Freon-13 boiling at -66ºC.

Expected net power: 2.6 kW.

Table 6C: Equipment list explanation for Table 6A

BLOWER Main air blower

COOLER fan aftercooler

CHILLER cooled air cooling device

SEPARAT Molecular Sieve Adsorption Air Dryer

HTEX main heat exchanger

EVAP External Cooling Evaporator

MEMBRANE Parallel membrane modules

HLEAK heat leak simulation

RCOMPR External cooling compressor

RCNDNS External cooling condenser

RJT External Cooling Expansion Valve

EXAMPLE 7 OXYGEN SYSTEM WITH MEMBRANE DEVICES IN CASCADES AND A RECYCLE FLOW

Example 7 is a process for producing an enriched oxygen permeate product stream from air. The oxygen stream specifications are the same as Example 5 except that the stream has an oxygen concentration of 85 percent on a molar basis. A simplified flow diagram for this process is shown in Figure 7. A summary of the major process streams is given in Table 7A and a summary description list of the major process equipment is given in Table 7B. Definitions of abbreviations used are given in Table 7C.

This procedure is similar to Example 5. However, here the permeate stream 740 from the first group of parallel membrane devices 733 is produced at medium pressure rather than low pressure, and the permeate stream 740 is used as feed to a second group of membrane devices 735. The low pressure permeate stream 741 from the second group of membrane devices 735 is the desired product, and the medium pressure non-permeate stream 760 is recycled through the main heat exchanger 727 to the inlet of the third stage of the four-stage compressor through a line 761. This cascading of membrane devices allows higher product oxygen concentrations to be achieved than with a single group of devices. The non-permeate stream 760 exiting the second group of membrane devices 735 contains 35 percent oxygen, so recycling the stream to the air compressor 705 helps to increase the oxygen content of the feed stream 731 to the first group of membrane devices 733, thereby increasing the driving forces for oxygen permeation through the membranes. Recirculating a medium pressure jet also helps to reduce energy consumption.

This example demonstrates how multiple membrane devices can be used in a cascade arrangement to achieve higher permeate product purities and how recycle streams can be used to increase the recovery and energy efficiency of a process.

See Table 7D for a summary of the simulation data generated by Examples 5-7 and described in Tables 5A through 7C. Table 7A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH CASCADED MEMBRANE DEVICES AND ONE RECYCLE STREAM Table 7A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH CASCADED MEMBRANE DEVICES AND ONE RECYCLE STREAM Table 7A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH CASCADED MEMBRANE DEVICES AND ONE RECYCLE STREAM Table 7A STREAM SUMMARY FOR A LOW TEMPERATURE OXYGEN PROCESS WITH CASCADED MEMBRANE DEVICES AND ONE RECYCLE STREAM

TABLE 7B LOW TEMPERATURE OXYGEN PROCESS WITH MEMBRANE MODULES IN CASCADE: Main Equipment Summary MAIN AIR COMPRESSOR

4-stage compressor with intercoolers and aftercooler.

Capacity: 0.45 kg/s air at 1545000 Pa delivery pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 201 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

MEMBRANE MODULES

First stage: 14.2 parallel modules containing 1.4 · 106 PBO-A fibers.

Second stage: 5.4 parallel modules containing 1.4 x 106 PBO-A fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 1.5748

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Brazed aluminum plate and fin heat exchanger with surface area expansion fins. Three groups of passages in counterflow arrangement.

Power, current 1: 0.69 kg/s at 1462000 Pa from 4.3ºC to -60ºC.

Stream 2: 0.21 kg/s at 122000 Pa from -153ºC to - 5.7ºC.

Stream 3: 0.25 kg/s at 336300 Pa from -60ºC to - 5.7ºC.

Useful power: 46.2 kW, maximum 10ºC warm end temperature difference.

Maximum pressure drop: 13800 Pa at any flow.

EXPANSION DEVICE

Generator-loaded turbo expansion device for extracting 23.1 kW of power.

Performance: 0.33 kg/s, 95 percent nitrogen at 1436000 Pa, -64ºC.

Pressure reduction to 122000 Pa.

Minimum efficiency: 80 percent isentropic

COOLER/DRYER

Heat exchanger, which is cooled by a nitrogen exhaust stream, and water separation tank.

Power, current 1: 0.70 kg/s at 1283000 Pa from 40ºC to 4.4ºC.

Flow 2: 0.33 kg/s at 114000 Pa from -60ºC to 15ºC.

Net power: 30.9 kW.

Maximum total pressure drop: 13800 Pa.

AIR DRYER

Double vessel molecular sieve adsorption air dryer with regenerative blower and heater.

Capacity: 0.70 kg/s saturated air at 1531000 Pa, 4.4ºC.

Dew point requirement: -70ºC at pressure.

Maximum pressure drop: 69000 Pa

OXYGEN COMPRESSOR

Centrifugal compressor with aftercooler.

Output: 0.10 kg/s 80% oxygen at 276000 Pa delivery pressure.

Minimum stage efficiency: 80 percent isentropic.

Engine power: 8.4 kW braking power.

Cooling capacity: maximum 40ºC outlet temperature at 20ºC ambient temperature.

Table 3C: Equipment list explanation for Table 7A

COMPRESS Main air compressor

AFT-COOL compressor aftercooler

CHILLER cooled air cooling device

SEPARATE Molecular Sieve Adsorption Air Dryers

HTEX main heat exchanger

MEMB1 First group of membrane modules

MEMB2 Second group of membrane modules

HLEAK heat leak simulation

EXPAND expansion device

N2SPLIT heat exchanger bypass divider

O2COMPR Product Oxygen Compressor TABLE 7D SUMMARY OF OXYGEN MEMBRANE METHODS

EXAMPLE 8 METHANE PURIFICATION PROCESS USING AN EXPANSION DEVICE FOR THE NON-PERMATE PRODUCT STREAM

Example 8 is a process for producing an enriched methane non-permeate product stream from methane-carbon dioxide mixtures. The feed gas stream is assumed to contain 95 percent methane and 5 percent carbon dioxide on a molar basis and is available at 1000 psia (6900 kPa) pressure and ambient temperature. A non-permeate product stream containing a minimum of 98 percent methane on a molar basis is desired with as high a recovery of the methane as possible. The enriched non-permeate product stream enters a pipeline for later compression to higher pressures, so it is desirable to minimize the overall pressure drop in the membrane process. This separation is often used in the natural gas industry where contaminated natural gas streams must be cleaned to meet pipeline specifications. A simplified flow diagram for a process for achieving this separation is shown in Figure 8. A summary of the major process streams is given in Table 8A, and the major process equipment is summarized in Table 8B. Definitions of abbreviations used are given in Table 8C.

It is assumed that the feed gas stream 830 is available at high pressure, therefore no compression equipment is shown. It is further assumed that the feed gas stream 830 has been scrubbed to remove substantially all particulate, condensable and corrosive components before entering the process. The scrubbed high pressure feed gas stream 830 then enters into the main heat exchanger 827 where it is cooled with heating product streams 840 and 852. The cooled feed gas stream 831 then enters the high pressure passages of the membrane device 833 and the carbon dioxide in the feed gas mixture preferentially permeates the membrane 834 to the low pressure passages of the membrane device 833. This causes the high pressure non-permeate product stream 850 to exit the membrane device 833 methane-enriched while the low pressure permeate product stream 840 exiting the membrane device is carbon dioxide-enriched. The high pressure non-permeate product stream 851 enters the expansion device 829 which removes some energy from the stream, lowering the pressure and temperature thereof as it exits the expansion device 829 via a line 852. This removal of energy offsets the heat entering the cold equipment through heat transfer with the environment and other sources, maintaining the equipment and process streams at the desired low temperature. The product streams 840 and 852 then enter the main heat exchanger 827 where they are heated with the cooling feed stream 830. The methane-enriched high pressure non-permeate product stream 853 then enters the product pipeline. The carbon dioxide-enriched low pressure permeate product stream 841 is typically passed to a burner where the heating value of the remaining methane in the permeate stream is restored.

This example demonstrates some of the features of a low temperature membrane process. Only a single membrane device is required to achieve the desired separation at given pressures. Due to the fact that the feed gas stream is available at much higher than ambient pressure, only the inlet pressure of the feed gas stream is required to provide the pressure difference at of the membrane. The cold process equipment and piping are well insulated to minimize heat leakage into the system. An expansion device is used to provide cooling and maintain the cold process equipment at the desired temperature. A heat exchanger is used to cool the feed stream and heat the product streams. By exchanging heat between these streams, the required cooling from the expansion device can be minimized, thereby minimizing the pressure drop suffered by the expansion device. No product recompression is necessary in this process since the methane-enriched non-permeate product enters a pipeline to an existing compressor. TABLE 8A STREAM SUMMARY FOR METHANE PURIFICATION PROCESS OF AN EXPANSION DEVICE ON THE NON-PERMATE STREAM TABLE 8A STREAM SUMMARY FOR METHANE PURIFICATION PROCESS OF AN EXPANSION DEVICE ON THE NON-PERMATE STREAM

TABLE 8B LOW TEMPERATURE METHANE PURIFICATION PROCESS USING AN EXPANSION DEVICE: MAIN EQUIPMENT SUMMARY MEMBRANE MODULES

1 module containing 1.4 · 10⁶ TCHF-BA-PC fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fibre effective length (m): 0.246

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Plate and fin heat exchanger with surface area expansion fins. Three groups of passes in counterflow arrangement.

Power, current 1: 2.2 kg/s at 6895000 Pa from 20ºC to -22ºC.

Stream 2: 2.0 kg/s at 6633000 Pa from -27ºC to 16ºC.

Flow 3: 0.2 kg/s at 158600 Pa from -22ºC to 16ºC.

Useful power: 260 kW, maximum 4.0ºC warm end temperature difference.

Maximum pressure drop: 20700 Pa at any flow.

EXPANSION DEVICE

Brake-loaded turbo expansion device for extracting 4.37 kW of power.

Output: 2.0 kg/s, 98 percent methane at 6818000 Pa, - 25ºC.

Pressure reduction to 6633000 Pa.

Minimum efficiency: 80 percent isentropic

Table 8C: Equipment list explanation for Table 8A

HTEX main heat exchanger

MEMBRANE Parallel membrane modules

HLEAK heat leak simulation

EXPAND expansion device

EXAMPLE 9 METHANE PURIFICATION PROCESS WITH AN EXTERNAL COOLING SYSTEM

Example 9 is also a process for producing an enriched methane non-permeate product stream from methane/carbon dioxide mixtures. The feed gas stream and required product specification for this example are the same as in Example 8. A simplified flow diagram for a process for carrying out this separation is shown in Figure 9. A summary of the major process streams is given in Table 9A and the major process equipment is summarized in Table 9B. The definitions of the abbreviations used are given in Table 9C.

This procedure is similar to that in Example 8 except that an external cooling system 965 is used in place of the expansion device to maintain the process equipment and process streams at the desired operating temperature. The feed gas stream 930 in this example is also assumed to be scrubbed and available at high pressure. The feed gas stream 930 enters a main heat exchanger 927 where it is cooled by the warming product streams 940 and 951. The feed gas stream 931 then enters an evaporative heat exchanger 964 where it is further cooled by a boiling coolant in line 962. The liquid coolant is provided by a simple single stage cooling system 965 which functions in the same manner as the external cooling system described in Example 3. However, since the process operates at a higher temperature than in Example 3, a different refrigerant is used and the refrigerant in line 960 exits the chiller compressor at a significantly higher temperature than in Example 3. The cold feed gas stream 932 exiting the evaporator 964 then enters the high pressure passages of the membrane device 933, and the carbon dioxide from the feed preferentially permeates through the membrane 934 to the low pressure passages of the membrane device 933. This leaves the non-permeate stream at high pressure 950 exiting the membrane device 933 enriched in methane, while the permeate stream at low pressure 940 exiting the membrane device 933 is enriched in carbon dioxide. The product streams 940 and 951 then enter the main heat exchanger 927 where they are cooled by the cooling feed gas stream 930. The methane-enriched permeate product stream at high pressure 952 then enters the product pipeline. The carbon dioxide-enriched permeate product stream at low pressure 941 is typically passed to a burner where the heating value of the residual methane in the permeate stream is recovered.

This process shows largely the same characteristics as Example 8. It also shows how the expansion device can be replaced by an external cooling system to maintain the process equipment at the desired operating temperature. TABLE 9A POWER SUMMARY FOR A METHANE PURIFICATION PROCESS WITH AN EXTERNAL COOLING SYSTEM TABLE 9A POWER SUMMARY FOR A METHANE PURIFICATION PROCESS WITH AN EXTERNAL COOLING SYSTEM TABLE 9A POWER SUMMARY FOR A METHANE PURIFICATION PROCESS WITH AN EXTERNAL COOLING SYSTEM

TABLE 9B LOW TEMPERATURE METHANE PURIFICATION PROCESS WITH EXTERNAL COOLING: MAIN EQUIPMENT SUMMARY MEMBRANE MODULES

1 module containing 1.4 · 10⁶ TCHF-BA-PC fibers.

Fiber ID (m): 9.5 · 10&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fiber effective length (m): 0.247

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Plate and fin heat exchanger with surface expansion fins. Three pass units in counterflow arrangement.

Power, current 1: 2.2 kg/s at 6895000 Pa from 20ºC to -20.7ºC.

Stream 2: 2.0 kg/s at 6804000 Pa from -25ºC to 16ºC.

Flow 3: 0.2 kg/s at 137900 Pa from -22ºC to 16ºC.

Net power: 251 kW, maximum 4.0ºC warm end temperature difference.

Maximum pressure loss: 20700 Pa per stream.

EXTERNAL COOLER COMPRESSOR

Freon-22 reciprocating compressor with the capacity 0.051 kg/s to compress from 150310 Pa to 1103000 Pa discharge pressure.

Engine power: 4.0 kW braking power.

EXTERNAL COOLING CONDENSER

Heat exchanger with the capacity to cool 0.051 kg/s of Freon-22 vapor at 1103000 Pa and 102ºC to 27ºC in ambient air at 20ºC.

Expected power: 12.2 kW.

EXTERNAL COOLER EVAPORATOR

Evaporative heat exchanger with the capacity to cool 2.2 kg/s methane at 6874000 Pa from -20.7ºC to -22ºC using boiling Freon-22 at -30ºC.

Expected power: 8.3 kW.

9C: Equipment list explanation for Table 9A

HTEX main heat exchanger

EVAP evaporator of the external chiller

MEMBRANE Parallel membrane modules

HLEAK heat loss simulation

RCOMPR Compressor of the external chiller

RCNDNS External chiller condenser

RJT Expansion valve of the external chiller

EXAMPLE 10 CARBON DIOXIDE RECOVERY PROCESS WITHOUT EXPANSION DEVICE OR EXTERNAL COOLING

Example 10 is a process for producing a carbon dioxide-enriched permeate product stream from methane/carbon dioxide mixtures. Similar separations are often encountered in the oil industry where carbon dioxide is recovered from the gas products in the enhanced residual oil recovery operation for re-injection. The feed gas stream is assumed to consist of 20 percent methane and 60 percent carbon dioxide on a molar basis and to be available at 500 psia (3,450 kPa) pressure and ambient temperature. A permeate product stream is required that has a minimum concentration of 92 percent carbon dioxide on a molar basis with as high a recovery of carbon dioxide as possible in the feed gas.

The carbon dioxide-enriched permeate product stream will enter a pipeline for later compression to higher pressures, so it is desirable to utilize any available energy from the process to compress the permeate product stream to an intermediate pressure. A simplified flow diagram for a process for carrying out this separation is shown in Figure 10. A summary of the major process streams is given in Table 10A, and the major process equipment is summarized in Table 10B. Definitions of the abbreviations used are given in Table 10C.

The feed gas stream 1030 is assumed to be available at high pressure, and therefore no compression equipment is shown. The feed gas stream is also assumed to be scrubbed, with all particulate components, condensable components and corrosive components removed prior to entering the process. Product gases from the enhanced residual oil recovery operation often contain higher order hydrocarbons than methane, which may be present at low temperatures. These are assumed to have been removed prior to entry into the process described herein. The scrubbed high pressure feed gas stream 1030 enters a main heat exchanger 1027 where it is cooled to -5°C by the warming high pressure non-permeate product stream 1051. The cooled feed gas stream 1031 enters the high pressure passages of the membrane device 1033 and the carbon dioxide from the feed preferentially permeates through the membrane 1034 to the low pressure passages of the membrane device 1033. This leaves the high pressure non-permeate stream 1050 exiting the membrane device 1033 enriched in methane while the low pressure permeate stream 1040 exiting the membrane device 1033 enriched in carbon dioxide. The methane-enriched permeate product stream at high pressure 1051 from the membrane device 1033 then enters the main heat exchanger 1027 where it is heated by the cooling feed gas stream 1030. The warm, carbon dioxide-enriched non-permeate product stream at high pressure 1052 enters the expansion device 1029 which removes some power from the stream, reducing the pressure to near ambient pressure. This methane-enriched non-permeate product stream 1053 is typically sent to a burner where the heating value of the residual methane is recovered. The cold, carbon dioxide-enriched permeate product stream 1040 from the membrane device 1033 is sent directly to a compressor 1074 which uses the line removed from the expansion device 1029 to compress the stream to an intermediate pressure. The carbon dioxide enriched permeate product stream 1041 exiting the compressor 1074 then enters a pipeline to which it is delivered for reuse in the enhanced recovery operation.

This process shows largely the same characteristics as Examples 8 and 9. However, here the desired product is the carbon dioxide-enriched permeate stream at low pressure rather than the methane-enriched non-permeate stream at high pressure.

It also shows how the expansion device can be replaced by an external cooling system to maintain the process equipment at the desired operating temperature. In addition, no process equipment is required to provide cooling to compensate for heat losses in the process. The expansion device is used only to take power from the product stream at high pressure to recompress the product stream at low pressure. No process equipment for cooling is required here because of the extent of the self-cooling ability of carbon dioxide. The Joule-Thomson effect on carbon dioxide results in a significant temperature reduction associated with the pressure loss as the carbon dioxide permeates through the membrane. The extent of the Joule-Thomson effect is sufficient to maintain the process equipment at the desired operating temperature.

Table 10D provides a summary of Examples 8-10. Examples 8-9 illustrate the extremely low methane loss in the carbon dioxide-enriched permeate stream. The BTU value of the permeate stream is sufficiently high to ensure burning of the permeate stream to recover the heating value. The amount of methane in the permeate stream, even if lost through flaring, is still less than the amount of methane required to regenerate the absorbent using conventional absorption technology to remove the carbon dioxide because conventional absorption technology uses as much as one cubic foot (CF) of methane per 1 CF of carbon dioxide removed. The methane is used to regenerate the absorbent. Example 10 illustrates a typical carbon dioxide recovery process for re-injection in an enhanced residual oil recovery operation. Example 10 shows that almost 97 percent of the carbon dioxide is recovered for re-injection. Example 10 illustrates the advantage of the low temperature process over ambient temperature carbon dioxide recovery processes which can only recover 50 percent of the carbon dioxide in the feed gas for re-injection. TABLE 10A POWER SUMMARY FOR A CARBON DIOXIDE RECOVERY PROCESS WITHOUT AN EXPANSION DEVICE TABLE 10A POWER SUMMARY FOR A CARBON DIOXIDE RECOVERY PROCESS WITHOUT EXPANSION DEVICE OR EXTERNAL COOLING SYSTEM

TABLE 10B LOW TEMPERATURE MEMBRANE CARBON DIOXIDE RECOVERY PROCESS: MAIN EQUIPMENT SUMMARY MEMBRANE MODULES

1 module containing 1.4 · 10⁶ TCHF-BA-PC fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fibre effective length (m): 1.27

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Plate and fin heat exchanger with surface expansion fins. Two pass units in counterflow arrangement.

Power, current 1: 4.14 kg/s at 3448000 Pa from 20ºC to -5ºC.

Stream 2: 0.86 kg/s at 3239000 Pa from -67ºC to 1.6ºC.

Useful power: 145 kW, maximum 4.0ºC warm end temperature difference.

Maximum pressure loss: 20700 Pa per stream.

EXPANSION DEVICE

Brake-loaded turbo expansion device for extracting 192 kW of power.

CO2 COMPRESSOR

Single-stage centrifugal compressor driven by the expander.

Performance: 3.28 kg/s 95 percent 002 at 726200 Pa discharge pressure.

Input conditions: 330960 Pa, -5ºC

Minimum stage efficiency: 80 percent isentropic.

EXAMPLE 11 LOW TEMPERATURE MEMBRANE PROCESS DEVELOPMENT: COMPARATIVE EXAMPLE

This example is designed to compare the yield and energy costs of nitrogen production between using a commercially available membrane material outside the scope of the present invention and using the low temperature membrane material PBO A of the present invention. Four process simulations were performed using commercially available process engineering simulation software. The first two processes use hollow fiber operated at 20°C. The third process uses PBO-A hollow fiber operated at -60°C using an expander in the non-permeate stream as described in Example 1. The fourth process uses PBO-A hollow fiber operated at -60°C using an external cooling system to cool the product gas stream to operating temperature as described in Example 3 without the main heat exchangers. The last procedure is included to illustrate the energy savings associated with the use of the main heat exchanger.

The calculations were performed on a membrane device containing 1,400,000 fibers with an effective length of 1.5748 meters (m) and a tubesheet thickness of 0.2032 m. The fibers are assumed to have an outside diameter of 1.35 x 10-9 m, an inside diameter of 9.5 x 10-5 m, and an internal separation area of 740 Angstroms thickness. The performance characteristics of the commercial membrane used for comparison were an oxygen/nitrogen separation factor of 7.4 and an oxygen permeability of 0.85. The desired gas product was specified to be present at a pressure of 125 psig (965 kPa). and not more than 1.0 percent oxygen. Table 11 lists the process recovery, yield of each membrane module, and energy costs to produce a given amount of product. This table illustrates that a module containing PBO-A membranes is significantly more efficient than a module containing the commercially available membranes outside the scope of the present invention and that the process described in Example 1 produces nitrogen at significantly lower energy costs. TABLE 10A POWER SUMMARY FOR A CARBON DIOXIDE RECOVERY PROCESS WITHOUT EXPANSION DEVICE OR EXTERNAL COOLING SYSTEM

TABLE 10B LOW TEMPERATURE MEMBRANE CARBON DIOXIDE RECOVERY PROCESS: MAIN EQUIPMENT SUMMARY MEMBRANE MODULES

1 module containing 1.4 · 10⁶ TCHF-BA-PC fibers.

Fiber ID (m): 9.5 x 10&supmin;&sup5;

Fiber OD (m): 1.35 x 10&supmin;&sup4;

Fiber density range thickness (m): 7.4 · 10⊃min;⊃8;

Fibre effective length (m): 1.27

Tube sheet length (m): 0.2032

MAIN HEAT EXCHANGER

Plate and fin heat exchanger with surface expansion fins. Two pass units in counterflow arrangement.

Power, current 1: 4.14 kg/s at 3448000 Pa from 20ºC to -5ºC.

Stream 2: 0.86 kg/s at 3239000 Pa from -67ºC to 1.6ºC.

Useful power: 145 kW, maximum 4.0ºC warm end temperature difference.

Maximum pressure loss: 20700 Pa per stream.

EXPANSION DEVICE

Brake-loaded turbo expansion device for extracting 192 kW of power.

CO2 COMPRESSOR

Single-stage centrifugal compressor driven by the expander.

Output: 3.28 kg/s 95 percent CO₂ at 726200 Pa discharge pressure.

Input conditions: 330960 Pa, -5ºC

Minimum stage efficiency: 80 percent isentropic.

EXAMPLE 11 LOW TEMPERATURE MEMBRANE PROCESS DEVELOPMENT: COMPARATIVE EXAMPLE

This example is designed to compare the yield and energy costs of nitrogen production between using a commercially available membrane material outside the scope of the present invention and using the low temperature membrane material PBO A of the present invention. Four process simulations were performed using commercially available process engineering simulation software. The first two processes use hollow fiber operated at 20°C. The third process uses PBO-A hollow fiber operated at -60°C using an expander in the non-permeate stream as described in Example 1. The fourth process uses PBO-A hollow fiber operated at -60°C using an external cooling system to cool the product gas stream to operating temperature as described in Example 3 without the main heat exchangers. The last procedure is included to illustrate the energy savings associated with the use of the main heat exchanger.

The calculations were performed on a membrane device containing 1,400,000 fibers with an effective length of 1.5748 meters (m) and a tubesheet thickness of 0.2032 m. The fibers are assumed to have an outside diameter of 1.35 x 10-9 m, an inside diameter of 9.5 x 10-5 m, and an internal separation area of 740 Angstroms thickness. The performance characteristics of the commercial membrane used for comparison were an oxygen/nitrogen separation factor of 7.4 and an oxygen permeability of 0.85. The desired gas product was specified to be present at a pressure of 125 psig (965 kPa). and not more than 1.0 percent oxygen. Table 11 lists the process recovery, yield of each membrane module, and energy costs to produce a given amount of product. This table illustrates that a module containing PBO-A membranes is significantly more efficient than a module containing the commercially available membranes outside the scope of the present invention and that the process described in Example 1 produces nitrogen at significantly lower energy costs.

Table 10C: Equipment list explanation for Table 10A

HTEX main heat exchanger

MEMBRANE Parallel membrane modules

HLEAK heat loss simulation

EXPAND expansion device

COMPRESS Main Air Compressor TABLE 10D SUMMARY OF CARBON DIOXIDE/METHANE MEMBRANE METHODS TABLE 11 Comparison of membrane processes

GENERAL PROCEDURE FOR PREPARING MEMBRANE SAMPLE PIECES AND EVALUATION IN EXAMPLES 12-22

Films for transport tests were generally prepared by preparing a solution of about 1 to about 10 percent of a polymer using the solvent indicated for each polymer, filtering the solution using a 5 micron filter, and pouring the solution onto a clean glass plate. The solvent was allowed to evaporate at a slow rate by partially covering the plate with a watch glass held a few millimeters above the surface of the solution. After the solvent had evaporated to dryness, the films were removed from the glass plate. The residual solvent was removed by placing the film in a vacuum oven at temperatures slightly below the softening temperature of the film (typically 160°C for polycarbonates and 300°C for the PBO materials) for at least 24 hours.

Circular test specimens were cut from the dried films. These specimens varied in size but generally ranged from 3 centimeters (cm) to 6 cm in diameter. The thickness of the film samples was measured using a film micrometer. These samples were then placed in cells which were sealingly attached to the film to create a feed and non-permeate chamber in contact with the first side of the membrane and a permeate chamber in contact with the second side of the membrane. The feed chamber was pressurized with a flowing stream of the feed gas mixture and the permeate chamber was maintained at near zero psi absolute (psia) by the test system. The feed chamber was designed in such a way as to ensure that all Areas of the membrane samples were exposed to the same feed gas mixture composition.

Analysis of the permeating gas stream consisted of determining both the identity and quantity of each gas in the stream using a mass spectrometer. Calibration of the spectrometer was performed periodically during the experiments by substituting a gas supply of known composition and flow rate for the unknown gas stream from the membrane cell. The feed gas stream was a binary mixture with one of the following compositions.

Feed gas composition (mole percent)

Oxygen 20

Nitrogen 80

Carbon dioxide 5

Methane 95

Helium-5

Methane 95

Feed gas pressures were generally maintained at about 20 to 45 psia (135 to 305 kPa) with a flow rate of about 25 cubic centimeters per minute (cc/min) through the feed chamber. Transport property measurements were made over a wide range of temperatures by controlling the temperature of the cell holder by circulating a heat transfer fluid for temperatures between -5ºC and 90ºC. Measurements at temperatures as deep as -30ºC were made by circulating liquid nitrogen through the coils in the cell holder. Measurements at even colder temperatures were made by attaching the permeation cell to the cold finger of a closed-loop variable temperature range helium-cooled cryostat (Janis Corp.) The results of the gas transport characterization measurements are expressed by the permeability of the highest permeability gas in units of Barrers, where 1 barrer = 1.0 x 10-10(cm3 (STP) cm)/(cm2 s cmHg) and a selectivity α which is defined as the ratio of the higher permeability gas to the lower permeability gas. The results for the examples are summarized in the tables below.

For each example, membrane characteristics of a polymer film were evaluated using the methods described above. Data for Examples 12 through 21 are given in Tables 12A through 22, respectively.

In Example 12, poly-4-methylpentene (PMP), available from Mitsui Petrochemical Industries Ltd. as "TPX Grade DX-810", was melt spun to form a 30 micrometer thick film, which was used as a membrane.

In Example 13, tetrachlorohexafluoro bisphenol A polycarbonate (TCHF BA PC) was evaluated as a membrane. The TCHF BA PC was prepared as follows. A 1.0 liter, four-necked, round-bottomed bottle equipped with a thermometer, condenser, phosgene/nitrogen inlet, and paddle stirrer connected to a Cole-Parmer Servodyne was charged with about 500 cubic centimeters of methylene chloride, 47.39 grams (0.1 mole) of tetrachlorohexafluoro bisphenol A, and 23.7 grams (0.3 mole) of pyridine. The resulting clear solution was stirred under a nitrogen atmosphere for ten minutes. Moderate stirring was continued, and about 9.9 grams (0.1 mole) of phosgene was added to the reaction mixture in a bubbling manner over a period of 27 minutes. The reaction mixture was then purged with methanol, with dilute hydrochloric acid and washed a second time with dilute hydrochloric acid. The solution was then passed through DOWEX-MSC-1 cation exchange resin (available from Dow Chemical Company) and the polymer was isolated by precipitation using hexane. The precipitated polymer was dried under vacuum at about 120°C for 48 hours. The resulting polycarbonate was found to have an inherent viscosity of about 0.29 dL/g at 25°C in methylene chloride. A homogeneous thin film was cast from a methylene chloride solution containing about 10 weight percent of the polymer.

In Example 14, tetrabromohexafluoro bisphenol A isophthalate ester (TBHF BA IE) was evaluated as a membrane. The TBHF BA IE was prepared as follows. A 1.0 liter, four-necked, round-bottom bottle equipped with a thermometer, condenser, phosgene/nitrogen inlet, and paddle stirrer connected to a Cole-Parmer servodyne was charged with about 500 cubic centimeters of methylene chloride, 41.68 grams (0.064 mole) of tetrabromohexafluoro bisphenol A, and 19.1 grams (0.241 mole) of pyridine. The resulting clear solution was stirred under a nitrogen atmosphere for ten minutes. Moderate stirring was continued, and about 12.98 grams (0.064 mole) of isophthaloyl chloride was added to the solution over a 20 minute period. The reaction mixture was then purified with methanol, neutralized with dilute hydrochloric acid and washed a second time with dilute hydrochloric acid. The solution was then passed through a DOWEX MSC-1 cation exchange resin and the polymer was isolated by precipitation using hexane. The precipitated polymer was dried under vacuum at about 120°C for 48 hours. The resulting polyisophthalate was found to have an inherent viscosity of about 0.16 dL/g at 25°C in methylene chloride. A homogeneous thin Film was cast from a methylene chloride solution containing about 10 weight percent of the polymer.

In Example 15, a polybenzoaxole having the formula PBO A was evaluated as a membrane. The PBO A was prepared as follows. A mixture of about 258 grams of 81 percent polyphosphoric acid (PPA) and 10.0 grams of diaminoresorcinol dihydrochloride (DAR) was added to a 500 milliliter resin kettle. The resin kettle was equipped with a nitrogen inlet, a silicone oil heat bath, a stainless steel stirrer shaft, and a high torque stirrer. The mixture was heated at about 110°C for 16 hours. At this time, about 89 grams of phosphorus pentoxide (P205) and 15.23 grams of 1,1,3-trimethyl-3-phenylindane-4,5'-dicarboxylic acid (PIDA) were added. The reactants were heated according to the following schedule: about 7 hours at 110°C, 16 hours at 150°C, 24 hours at 180°C, and 24 hours at 190°C. The crude polymer was isolated by precipitation in water, vacuum filtering, washing with hot water and methanol, and finally drying in a vacuum oven. The polymer was soluble in m-cresol, trifluoroacetic acid, and methanesulfonic acid. A homogeneous thin film was cast from a m-cresol solution containing about 8 weight percent of the polymer.

In Example 16, a polyoxazole having the formula of PBO B was tested. The PBO B was prepared as follows. A mixture of about 55 grams of 81 percent polyphosphoric acid, 5.33 grams of 2,2-bis(3-amino-4-hydroxyphenyl)-1,1,1,3,3,3-hexafluoropropane, 20 grams of phosphorus pentoxide (P₂O₅), and 3.76 grams of diphenyl ether-4,4-dicarboxylic acid was added to a 500 ml resin kettle. The resin kettle was equipped with a nitrogen inlet, a silicone oil heat bath, a stainless steel stirrer shaft, and a high torque stirrer. The reactants were heated according to the following schedule: about 2 hours at 70°C, 3 hours at 140ºC, 2 hours at 160ºC, 8 hours at 180ºC and 3 hours at 200ºC. The crude polymer was isolated by precipitation in water, washing with hot water and methanol and finally drying in a vacuum oven. The polymer was soluble in m-cresol, trifluoroacetic acid and methanesulfonic acid. A homogeneous thin film was cast from a m-cresol solution containing about 10% by weight of the polymer.

In Example 17, polyphenylene oxide (PPO) was tested as a membrane. The PPO was obtained as a commercial polymer from Aldrich Chemical Company. A film about 25 micrometers thick was formed by compression molding at about 300°C.

In Example 18, a polymer formed by the condensation of hexafluoro bisphenol A and decafluoro bisphenyl (HFBA/DFBP) was investigated as a membrane. The HFBA/DFBP was synthesized by the following procedure. Hexafluoro bisphenol A, about 25.15 grams (0.748 mole), decafluoro bisphenyl, about 25.00 grams (0.0748 mole), potassium carbonate, about 41.37 grams (0.299 mole), and dimethylacetamide, 275 milliliters, were added to a 500 milliliter round-bottom three-necked flask equipped with a mechanical stirrer, a thermometer, and a nitrogen inlet. The mixture was stirred and heated to about 60°C under a nitrogen atmosphere for about 12 hours. Afterward, the mixture was poured into 6 liters of water with stirring. The white solid which formed was isolated by filtration and left in 3 liters of water overnight. The solid was then isolated by filtration. The solid was dissolved in about 400 milliliters of methylene chloride, dried with magnesium sulfate, filtered through a filter aid, and precipitated with stirring into 2 liters of methanol. The solid was isolated by filtration and dried in a vacuum oven at about 30°C to give about 36.00 grams of a white solid, yielding of 76 percent. The inherent viscosity of the material in chloroform was determined to be about 1.064. Differential scanning calorimetry (DSC) analysis indicated a glass transition temperature of about 180°C. Proton NMR analysis of the product was consistent with the structure of the desired product. A membrane was prepared by casting a film from a solution containing about 10 weight percent of the polymer in methylene chloride.

In Example 19, an indane-containing polyimide obtained from the condensation of pyromellitic dianhydride and 1,1,3-trimethyl-3-phenylindane-4,5-diamine (PMDA/PIDA) was evaluated as a membrane. To form the monomer, about 100 grams (0.308 moles) of 1,1,3-trimethyl-3-phenylindane-4,5'-dicarboxylic acid, available from Amoco, was added to about 375 milliliters of concentrated sulfuric acid in a 4-liter beaker. About 2,500 milliliters of chloroform were added to the beaker to form a two-phase mixture. The mixture was stirred at about 40°C to 45°C. Over a period of about 45 minutes, sodium azide, about 48.8 grams (0.751 moles), was slowly added. The temperature was maintained below about 45°C during the addition of sodium azide. Thereafter, the temperature was raised to about 55°C for about one hour, and the mixture was then cooled to room temperature. The solution was basified by the careful addition of sodium hydroxide, which caused some of the diamine to precipitate as a paste. The paste was collected, dissolved in methylene chloride, and combined with the organic portion. The organic solutions were washed with water and dried with anhydrous potassium carbonate. The solvent was removed by rotary evaporation, and the resulting semi-solid was crystallized from an ether/hexane mixture. The monomer product yield of about 52.5 grams was about 64 percent. The monomer product had a melting point of about 94°C to 95°C. To form the polymer, about 2,000 grams (7.508 mmol) of 5-amino-1,1,3-trimethyl-3-(4-aminophenyl)indane, about 1.638 grams (7.508 mmol) of 1,2,4,5-benzenetetracarboxylic acid dianhydride, about 0.75 grams (5.8 mmol) of isoquinoline, and about 36 milliliters of m-cresol were added to a 100 milliliter round-bottom bottle. The bottle was fitted with a condenser, sparged with dry nitrogen, and stirred with a magnetic stir bar while heating to reflux for one hour at about 200°C. The solution was cooled to about 80°C and the viscous solution was poured into about 200 milliliters of methanol. After breaking up the fibrous solids in a blender, the resulting yellow solid was washed with methanol and dried under vacuum at about 100°C for about 12 hours. An about 5 percent solution of the indane-containing polyimide in hot m-cresol was prepared. The solution was poured onto a clean glass plate using a 20-mil die cutter. The solvent was removed from the film under a nitrogen purge in an oven at about 100°C for a period of about three hours. The plate and film were removed from the oven and allowed to cool. The film was removed from the plate by immersion in water. The film was dried under vacuum at about 300°C for about 30 minutes.

In Example 20, a polymer obtained from the condensation of 1,4-diaminobenzene-2,5-dithiol dihydrochloride and 1,1,3-trimethyl-3-phenylindane-4',5-dicarboxylic acid (DABT/PIDA) was evaluated as a membrane. DABT/PIDA was prepared by adding about 5.00 grams (0.0204 moles) of 1,4-diaminobenzene-2,5-dithiol dihydrochloride, about 6.615 grams (0.0304 moles) of 1,1,3-trimethyl-3-phenylindane-4',5-dicarboxylic acid (PIDA), and about 184 grams of 85% phosphoric acid to a 1 liter resin kettle. The vessel was continuously sparged with nitrogen and an air driven stainless steel stirrer was used for mixing. The reaction mixture was heated according to the following schedule: about 16 hours at about 50°C, about 8 hours at about 90°C, and about 20 hours at about 110°C. About 8 grams of phosphorus pentoxide (P₂O₅) was added and the mixture heated at about 135°C for about 24 hours, about 150°C for about 8 hours, and finally at about 190°C for about 24 hours. The extremely viscous solution was coagulated using a mixer and 2 liters of water, filtered, washed with water, washed with methanol, and dried under vacuum at about 150°C. This yielded about 8.5 grams of a light green polymer having an inherent viscosity of about 18 dL/gram in methanesulfonic acid/methanesulfonic anhydride. To prepare a film from the polymer, about 0.5 grams of the polymer and 12 grams of m-cresol were placed in a small glass vial. The contents of the vial were heated and stirred to dissolve the polymer in the solvent. The viscous brown solution was pressure filtered through a 10 micron Teflon filter; the solution was poured onto a clean glass plate and placed in a 50°C oven under a nitrogen stream for about 3 hours. A vacuum was applied to the oven for about 2 hours and then the temperature was raised to about 110°C for about two hours. The cooled plate and film were placed in a tray of distilled water to aid in the removal of the film from the plate. The film was dried under vacuum at about 300ºC for about 1 hour.

In Example 21, 9,9-bis(3,5-dibromo-4-hydroxyphenyl)fluorene polycarbonate (TBF PC) was evaluated as the membrane. The TBF PC was prepared as follows. A three-necked 0.5 liter round bottom bottle fitted with a condenser, a phosgene/nitrogen inlet and a Cole-Parmer Servodyne was charged with about 237 cubic centimeters of methylene chloride, 30.80 grams (0.046 moles) of 9,9-bis(3,5-dibromo-4-hydroxyphenylfluorene, and 10.9 grams (0.138 moles) of pyridine. The resulting clear solution was stirred under a nitrogen atmosphere, and about 4.6 grams (0.046 moles) of phosgene was added to the reaction mixture in a bubbling manner over a period of about 7 minutes. An additional amount of about 1.0 gram (0.010 moles) of phosgene was added in a bubbling manner over about 18 minutes, and the reaction mixture was stirred for about 16 hours. Thereafter, the reaction mixture was purified with methanol, diluted with about 50 cubic centimeters of methylene chloride, washed twice with dilute hydrochloric acid, and then passed through DOWEX MSC-1 ion exchange resin. The polymer was passed through Adding the methylene chloride solution of the polymer to a mixture of hexane/acetone. The precipitated polymer was dried under vacuum at about 120°C for about 48 hours. The resulting polycarbonate was found to have an inherent viscosity of about 0.48 dL/g at 25°C in methylene chloride. A homogeneous thin film was cast from a methylene chloride solution containing about 10 weight percent of the polymer (identified as TBF PC No. 1 in Table 21).

A large scale preparation of TBF PC was carried out as follows. A 2.0 liter, four-necked, round bottom bottle equipped with a thermometer, a condenser, a phosgene/nitrogen inlet, and a paddle stirrer connected to a Cole-Parmer Servodyne was charged with about 1.0 liter of methylene chloride, 129.95 grams (0.194 moles) of 9,9-bis(3,5-dibromo-4-hydroxyphenyl)fluorene, and 46.0 grams (0.582 moles) of pyridine. The resulting clear solution was stirred under a nitrogen atmosphere, and about 20.2 grams (0.204 moles) of phosgene was added over a period of about 36 minutes in a bubbling manner to the reaction mixture. The reaction mixture was stirred for about 18 hours, 1.0 gram of 9,9-bis(3,5-dibromo-4-hydroxyphenyl)fluorene was added, the reaction mixture was stirred for about 50 minutes, and about 0.1 gram of phosgene was added. The reaction mixture was stirred for about 30 minutes, was diluted with about 200 cubic centimeters of methylene chloride and then purified with methanol. The reaction mixture was washed twice with dilute hydrochloric acid and then passed through DOWEX MSC-1 ion exchange resin. The polymer was isolated by adding the methylene chloride solution of the polymer to a mixture of hexane/acetone. The precipitated polymer was dried under vacuum at about 120°C for about 72 hours. The resulting polycarbonate was found to have an inherent viscosity of about 0.81 dL/g at 25°C in methylene chloride. A homogeneous thin film was cast from a methylene chloride solution containing about 10 weight percent of the polymer (identified as TBF PC No. 2 in Table 21).

In Example 22, 9,9-bis(3,5-dichloro-4-hydroxyphenyl)fluorenecarbonate (TCF PC) was evaluated as a membrane. The TCF PC was prepared as follows. A three-necked 0.5 liter round bottom bottle equipped with a condenser, a phosgene/nitrogen inlet, and a paddle stirrer connected to a Cole-Parmer Servodyne was charged with about 150 cubic centimeters of methylene chloride, 19.00 grams (0.039 moles) of 9,9-bis(3,5-dichloro-4-hydroxyphenyl)fluorene, and 9.25 grams (0.117 moles) of pyridine. The resulting clear solution was stirred under a nitrogen atmosphere, and about 4.0 grams (0.040 moles) of phosgene was added to the reaction mixture in a bubbling manner over a period of about 12 minutes. The reaction mixture was stirred for about 1 hour and was then purged with methanol. The reaction mixture was washed twice with dilute hydrochloric acid, then with water, and through DOWEX MSC-1 ion exchange resin. The polymer was isolated by adding the methylene chloride solution of the polymer to a mixture of hexane/acetone. The precipitated polymer was dried under vacuum at about 120°C for about 48 hours. The resulting polycarbonate was found to have an inherent viscosity of about 0.55 dL/g at 25°C in methylene chloride. A homogeneous thin film was cast from a methylene chloride solution comprising about 10 weight percent of the polymer.

The results for each example are tabulated below in tables identified by the example number. Although relatively low transmembrane pressures were used in most of the examples, evaluations at a transmembrane pressure of 165 psia (1140 kPa) produced results similar to those at a lower pressure of oxygen/nitrogen. TABLE 12A EXAMPLE 12 - PMP TABLE 13A EXAMPLE 13 - TCHF-BA-PC TABLE 13A EXAMPLE 13 - TCHF-BA-PC TABLE 13B EXAMPLE 13 - TCHF-BA-PC TABLE 13B EXAMPLE 13 - TCHF-BA-PC TABLE 13C EXAMPLE 13 - TCHF-BA-PC TABLE 13C EXAMPLE 13 - TCHF-BA-PC TABLE 14A EXAMPLE 14 - TBHF-BA-IE TABLE 15A EXAMPLE 15 - PBO-A TABLE 15B EXAMPLE 15 - PBO-A TABLE 15B EXAMPLE 15 - PBO-A TABLE 15C EXAMPLE 15 - PBO-A TABLE 15C EXAMPLE 15 - PBO-A TABLE 16A EXAMPLE 16 - PBO-B TABLE 16B EXAMPLE 16 - PBO-B TABLE 16B EXAMPLE 16 - PBO-B TABLE 16C EXAMPLE 16 - PBO-B TABLE 17A EXAMPLE 17 - PBO TABLE 18C EXAMPLE 18 - HFBA/DFBP TABLE 19A EXAMPLE 19 - PMDA/PIDA TABLE 19B EXAMPLE 19 - PMDA/PIDA TABLE 19C EXAMPLE 19 - PMDA/PIDA TABLE 20B EXAMPLE 20 - DABT/PIDA TABLE 20C EXAMPLE 20 - DABT/PIDA TABLE 21A EXAMPLE 21 - TBF-PC #1 TABLE 21A EXAMPLE 21 - TBF-PC #1 TABLE 21A SAMPLE 21 - TBF-PC #2 TABLE 22A EXAMPLE 21 - TBF-PC #2

EXAMPLE 23 MIXED GAS TEMPERATURE AND PRESSURE INFLUENCES ON TCHF-BA-PC

The TCHF-BA-PC membrane of Example 13 was used to evaluate the effects of temperature and pressure on membrane performance using mixed gas feeds. The apparatus described above was used to evaluate membrane performance at different temperatures and pressures using a feed gas mixture of 2 mole percent hydrogen and 98 mole percent methane, a mixture of 50 mole percent hydrogen and 50 mole percent methane, a mixture of 2 mole percent hydrogen and 98 mole percent ethane, and a mixture of 50 mole percent CO2 and 50 mole percent methane. The data are plotted in Tables 23A-C. In the hydrogen/hydrocarbon experiments, the permeation rate of hydrogen almost immediately reached the steady-state rate compared to the permeation rate of the hydrocarbon gases. Such behavior allows an estimate of the clean gas permeabilities to be made. In almost all cases, the presence of hydrocarbon gas resulted in a decrease in hydrogen permeability compared to the estimated clean gas values. This effect was more pronounced at lower temperatures, higher pressures and/or for hydrocarbon gases that were more soluble in the membrane material.

In the carbon dioxide/methane experiments, carbon dioxide softened the membrane, producing a loss of selectivity that increased at lower temperatures and/or higher pressures. TABLE 23A1 EXAMPLE 23 - TCHF-BA-PC (feed = 2 mole percent H₂ and 98 mole percent CH₄) TABLE 23A2 EXAMPLE 23 - TCHF-BA-PC (feed = 50 mole percent H₂ and 50 mole percent CH₄) TABLE 23B EXAMPLE 23 - TCHF-BA-PC TABLE 23C EXAMPLE 23 - TCHF-BA-PC

EXAMPLE 24 MIXED GAS TEMPERATURE AND PRESSURE INFLUENCES ON TBF-PC

The TBF-PC membranes in Example 21 were tested at higher and lower feed pressures at -42ºC to evaluate potential for mixed gas effects under elevated pressure and low temperatures. A feed mixture of 80 mole percent nitrogen and 20 mole percent oxygen was used. The data are plotted in Table 24.

When the feed pressure was increased from 2.5 to 157 psi, the permeabilities of oxygen and nitrogen decreased by about 17 percent, while the selectivity remained essentially unchanged.

Single gas oxygen and nitrogen permeability measurements were made at approximately -42ºC using an oxygen feed pressure of 20.0 psi and a nitrogen feed pressure of 124.0 psi. The oxygen permeability was 2.08 Barrers and the nitrogen permeability was 0.130 Barrers. The oxygen/nitrogen selectivity in these single gas measurements was 16.0. The single gas selectivity is approximately 20 percent higher than the mixed gas selectivity.

EXAMPLE 25 QUALITY NUMBER

Plotted in Table 25 are the values of the figures of merit for oxygen/nitrogen, carbon dioxide/methane and helium/methane separation as calculated using Equations III, IV and V as previously defined herein for membranes prepared with the following polymers: a polybenzoxazole (PBO A), the condensation polymer of pyromellitic dianhydride and 1,1,3-trimethyl-3-phenylindane-4',5-diamine (PMDA/PIDA), poly-4-methylpentene-1 (PMP), tetrachlorohexafluorobisphenol-A polycarbonate (TCHF-BA-PC), a copolymer of 50 percent tetrabromobisphenol-A polycarbonate and 50 percent hexafluorobisphenol-A (TBBA/HFBA), polyphenylene oxide (PPO), a polybenzoxazole (PBO-B), tetrabromohexafluorobisphenol-A polycarbonate (TBHF-BA-PC), the condensation polymer of Hexafluorobisphenol-A and decafluorobisphenyl (HFBA/DFBP), the condensation polymer of 1,4-diaminobenzene-2,5-dithiol and 1,1,3-trimethyl-3-phenylindane-4',5-dicarboxylic acid (DABT/PIDA), 9,9-bi(3,5-dibromo-4-hydroxyphenyl)fluorene polycarbonate (TBF-PC), tetrabromobisphenol-A polycarbonate (TBBA-PC), TRYCITE polystyrene (PST), cellulose triacetate (CTA) and unsaturated polysulfone (PSF). The data for TBBA-PC, CTA, PST and PSF are shown for comparison purposes only and are not illustrative of the present invention. TABLE 25 QUALITY NUMBER¹

as calculated using equations III, IV and V at 30ºC.

EXAMPLE 26 PERMEATION DATA

The properties of the unsaturated fluorene polycarbonate (BHPF-PC) measured by Paul and Aguilar (University of Texas, April 14, 1992) at 35ºC are plotted in the row of Table 24 in the non-brominated section. Our measurements on both BHPF-PC and TBBHPF-PC are plotted in the fourth row. Thus, these values for TBBHPF-PC (brominated) show a significant increase in both permeability and selectivity, rather than an increase in selectivity and a decrease in permeability as described in the literature. TABLE 24

BA-PC = (Bi-phenol A polycarbonate)

BHPF-PC = (polymer from TBF-PC above)

Considering these factors, there is increased commercial value in the transport properties of TBBHPF-PC (brominated) at ambient temperature, 20ºC and below, such as at temperatures 1, 2, 3, 4 or 5ºC above freezing, and (2) there is an unexpected increase in both O2 permeability and O2/N2 selectivity due to bromination.

Claims (56)

1. A process for separating component gases in a gas mixture, comprising: A/ Contacting a first side of a gas separation membrane, the membrane comprising a polymeric separation layer or region which is in a glassy state under the conditions in which the separation is carried out, with a gas mixture while maintaining a difference in chemical potential from the first side of the membrane to a second side of the membrane such that at least a first component gas of the gas mixture selectively permeates to at least a second component gas in the gas mixture from the first side of the membrane to the second side of the membrane, said contacting occurring at a temperature of 5°C or below, provided that the permeation occurs at or above the freezing point of any liquid present, the membrane being selected so that when using a mixture of 80 mole percent nitrogen and 20 mole percent oxygen as feed at 30°C at a pressure of 205.5 kPa (30 psia) on the first side of the membrane and a vacuum of less than 0.133 kPa (1 mm Hg) on the second side of the membrane, the permeability of oxygen in Barrers is less than 2000 and has the following relationship for oxygen/nitrogen selectivity: Permeability > 2000/(selectivity)7/2; and B/ Recovering at least one of the permeate gas or the directed non-permeate gas, wherein said gas separation effecting layer/region is a glassy polymeric composition of material having a repeating unit of the formula: where U is independently selected from chlorine or bromine. 2. The method of claim 1, wherein the membrane is an asymmetric membrane, anisotropic membrane or composite membrane. 3. A method as described in any one of claims 1 or 2, wherein the membrane is in the form of a plurality of hollow fibers. 4. A method according to claim 3, wherein the gas mixture is introduced into the bores of the hollow fibers and the gas which has permeated through the fibers flows generally countercurrently to the gas in the bores of the hollow fibers. 5. A method as described in any one of claims 1 or 2, wherein the membrane is in the form of a flat sheet or film. 6. A method as described in claim 3 or 4, wherein the hollow fibers form a separation layer or a separation region in the near the inner surface of the hollow fibers and a generally porous region extending from the separation layer or region to the outer surface of the hollow fibers. 7. A method as described in any one of claims 3 or 4, wherein the hollow fibers comprise a generally porous region at or near the outer surface of the hollow fibers, a generally porous region at or near the inner surface of the hollow fibers, and a separating layer or region located generally between the two porous surfaces. 8. A process according to any preceding claim, wherein the first component gas is oxygen and the second component gas is nitrogen. 9. A process as described in any one of claims 1 to 7, wherein the first component gas is carbon dioxide and the second component gas is light hydrocarbon. 10. A process as described in any one of claims 1 to 7, wherein the first component gas is hydrogen or helium and the second component gas is a light hydrocarbon. 11. A process as described in any one of claims 1 to 7, wherein the first component gas is ethylene and the second component gas is ethane. 12. A process as described in any one of claims 1 to 7, wherein the temperature of separation is in the range of -30°C to -5°C. 13. A process as described in claim 8, wherein the temperature of separation is in the range of -150°C to -5°C. 14. A process as described in claim 12, wherein the temperature of separation is in the range of -125°C to -25°C. 15. A process as described in claim 9, wherein the temperature of separation is in the range of -40°C to -5°C. 16. A process as described in any one of claims 1 to 7, wherein the membrane has a permselectivity index of at least 0.25 for oxygen/nitrogen separation. 17. A process as described in any one of claims 1 to 7, wherein the membrane has a figure of merit for the separation oxygen/nitrogen of at least 3.0. 18. A process as described in any one of claims 1 to 7, wherein the membrane has a figure of merit for the separation of carbon dioxide/methane of at least 25. 19. A process as described in any one of claims 1 to 7, wherein the membrane has a figure of merit for helium/methane separation of at least 15. 20. A process as described in any one of claims 1 to 7, wherein the membrane is at a temperature of at least -5°C or less has an oxygen/nitrogen selectivity of at least 12 and an oxygen permeability of at least 1 Barrer as measured using a mixture of 80 mole percent nitrogen and 20 mole percent oxygen as feed at a pressure of 205.5 kPa (30 psia) on the first side of the membrane and a vacuum of less than 0.133 kPa (1 mm Hg) on the second side of the membrane. 21. A process as described in any one of claims 1 to 7, wherein the glassy polymer has a glass transition temperature of at least 100°C. 22. A process as described in any one of claims 1 to 7, wherein the membrane has a figure of merit for the separation oxygen/nitrogen of at least 1.0. 23. A process as described in any one of claims 1 to 7, wherein the membrane has a figure of merit for the separation carbon dioxide/methane of at least 20. 24. A process as described in any one of claims 1 to 7, wherein the membrane has a carbon dioxide/methane selectivity of at least 40 and a carbon dioxide permeability of at least 10 Barrers at a temperature of 5°C or less, as measured using a mixture of 5 mole percent carbon dioxide and 95 mole percent methane as feed with a pressure of 205.5 kPa (30 psia) on the first side of the membrane and a vacuum of less than 0.133 kPa (1 mm Hg) on the second side of the membrane. 25. A method as described in any one of claims 1 to 7 wherein the membrane has a figure of merit for the separation of helium/methane of at least 5. 26. A process as described in any preceding claim, further comprising recovering at least one rejected non-permeate gas from the first side of the membrane and at least one permeate gas from the second side of the membrane. 27. A process as described in any preceding claim, wherein the temperature of the gas mixture is changed by cooling the gas mixture using at least one of an external cooling system, an absorption cooling system, an expansion device or the cooling capacity of the permeate gas. 28. A process as described in any preceding claim, comprising the additional step of expanding at least a portion of either the feed gas mixture, the non-permeate gas or the permeate gas to cool the gas. 29. A process as described in claim 27, wherein the gas is expanded and cooled using a turbine expander. 30. A method according to any one of the preceding claims, comprising the additional step of passing the cooled gas through a device for exchanging heat with the gas mixture to be introduced into the membrane to cool the gas mixture. 31. A method according to any preceding claim, comprising the additional step of pretreating the gas mixture to substantially remove one or more impurities which have a detrimental effect on the performance or physical integrity of the membrane or system components. 32. A process as described in claim 31, wherein the contaminants comprise solid particles or liquid droplets larger than 1 micron (um), hydrocarbons, glycols, oil field additives, surfactants, mercury, oils, hydrogen sulfide, water vapor, nitrogen oxides, sulfur oxides, hydrogen chloride, chlorine, mercaptans, carbon disulfide, or a combination thereof. 33. A process as described in any preceding claim comprising the additional step of recycling at least one of non-permeate gas or permeate gas. 34. A process as described in claim 30 comprising the additional step of using at least a portion of either the non-permeate gas or the permeate gas to purge or regenerate pretreatment equipment. 35. A process as described in any preceding claim operated with two or more membrane devices arranged in series or cascade or a combination thereof. 36. A method as described in claim 35, comprising the additional step of compressing gas between the devices. 37. A method as described in claim 35, wherein the composition of the membrane used in at least one device is different from the composition of the membrane used in at least one other device. 38. A method as described in claim 35, wherein the temperature of one device is different from the temperature of at least one other device. 39. A process as described in any preceding claim, comprising the additional separation step of at least one of pressure swing adsorption, temperature swing adsorption, vacuum swing adsorption, fractionalization, distillation, rectification, chemical adsorption, chemical reaction or extraction. 40. A method as described in claim 33, wherein the recycled gas is used as a purge gas on the second side of the membrane. 41. Device for separating components of a gas mixture, comprising: means (127) for changing the temperature of the gas mixture to a temperature of 5ºC or lower and at or above the freezing point of any liquid present, the means for changing the temperature having an inlet and an outlet; at least one membrane device (133) operating at said temperature or less and having an inlet for introducing a gas mixture to be separated, a membrane (134) which divides the membrane device into a first side of the non-permeate chamber of the membrane and a second side of the permeate chamber of the membrane, has an outlet for non-permeate gas and an outlet for permeate gas, means (130) for conveying gas from the outlet of the means (127) for changing the temperature of the gas mixture to the inlet of the membrane means (133), wherein the membrane within the membrane means is as defined in any one of claims 1 to 25. 42. Apparatus according to claim 41, further comprising means (105) for compressing having an inlet and an outlet, and conveying means for conveying the gas from the outlet of the compression means to the inlet of the change means. 43. Apparatus for separating the components of a gas mixture according to claim 42, wherein said means for changing the temperature of the gas mixture comprises: at least one heat exchanger device (127) having a first inlet and a first outlet and a second inlet and a second outlet, a conduit connecting the first inlet and the first outlet and a conduit connecting the second inlet and the second outlet, said first and second conduits being in thermal contact along a major portion of their respective lengths, said means for conveying the gas mixture from the outlet of the gas compression device (105) to the inlet of said temperature changing device for the gas Means (130) for conveying the gas mixture from the outlet of the means (105) for compressing the gas to the first inlet for the heat exchanger means (127), said means for conveying gas from the outlet of the means for changing the temperature of the gas to the inlet of the membrane means (133) is means (132) for conveying gas from the first outlet of the heat exchanger means (127) to the inlet of the membrane means (133), and further comprising additional means for changing the temperature of a portion or all of the gas exiting either the permeate gas outlet or the non-permeate gas outlet of the membrane means (133), or a combination thereof, said additional means (129) for changing the temperature of the gas having an inlet and an outlet, means (150) for conveying a portion or all of the gas from either the permeate gas outlet or the non-permeate gas outlet of the membrane means (133) or a combination thereof to the inlet of said additional means (129) for varying the temperature of the gas, and means (15) for conveying gas from the outlet of said additional means for changing the temperature of the gas to the second inlet of the heat exchanger means (127). 44. Apparatus for separating components of a gas mixture according to claim 42, wherein said means for changing the temperature of the gas mixture comprises: at least one heat exchanger device (127) having a first inlet and a first outlet and a second inlet and a second outlet, a first inlet and the first outlet and a line connecting the second inlet and the second outlet, these first and second lines being in thermal contact along a major portion of their respective lengths, said means for conveying the gas mixture from the outlet of the gas compression device (105) to the inlet of said temperature changing device for the gas is a device (130) for conveying the gas mixture from the outlet of the device (105) for compressing the gas to the first inlet for the heat exchanger device (127), and further comprising an additional device (128) for changing the temperature of a part or all of the gas, this additional device for changing the temperature of the gas having an inlet and an outlet, a device (131) for conveying the gas mixture from the first outlet of this heat exchanger device (127) to the inlet of this additional device (128) for changing the temperature of the gas, means (132) for conveying gas from the outlet of this additional means for changing the temperature of the gas to the inlet of the membrane means (133), and means (150-152) for conveying a portion or all of the gas from either the permeate gas outlet or the non-permeate gas outlet of the membrane means or a combination thereof to the second inlet of the heat exchanger means (127). 45. Apparatus as described in any one of claims 41 to 43, wherein the means for changing the temperature of the gas is an external cooling system. 46. Apparatus as described in any one of claims 41 to 43, wherein the means for changing the temperature of the gas is an adsorption cooling system. 47. Apparatus as described in any one of claims 41 to 43, wherein the means for changing the temperature of the gas is an expansion device. 48. . Device as described in any one of claims 41 to 43, further comprising: a device (126) for pretreating the gas mixture to substantially remove contaminants that have a detrimental effect on the physical integrity and/or the performance of the membrane, the device for pretreating the gas having an inlet and an outlet, and a device (130) for conveying the pretreated gas to the inlet of the device (105) for compressing the gas, the inlet for changing the temperature of the gas, the first inlet of the heat exchanger device (127) or the inlet of the membrane device (133). 49. Apparatus as described in any one of claims 41 to 43, wherein two or more membrane devices (733, 735) are arranged in series or cascade arrangement or a combination thereof. 50. Apparatus as described in claim 49, wherein the composition of the membrane used in at least one stage is different from the composition of the membrane used in at least one other stage. 51. Device as described in any one of claims 41 to 43, additionally comprising: means for recycling a portion or all of the gas from either the permeate gas outlet or the non-permeate gas outlet of the membrane means or a combination thereof to the inlet of the means for compressing the gas, the inlet for changing the temperature of the gas, the first inlet of the heat exchanger means or the inlet of the membrane means. 52. Apparatus as described in claim 51, wherein the membrane means has an inlet for purge gas and the apparatus additionally comprises: means for conveying a portion or all of the gas from either the permeate gas outlet or the non-permeate gas outlet of the membrane means, or a combination thereof, to the purge gas inlet of the membrane means. 53. Apparatus as described in claim 49, which additionally comprises: means for compressing a portion or all of the gas from either the permeate gas outlet or the non-permeate gas outlet of the membrane means or a combination thereof. 54. A device for separating components of a gas mixture according to claim 41, further comprising: a device for recovering the compression energy of some or all of the gas from either the permeate gas outlet or the non-permeate gas outlet of the membrane device or a combination thereof, the recovered energy being used to drive a compression device or to generate electrical power. 55. Apparatus for carrying out the separation process of claim 1, comprising: a device for compressing the gas mixture to a desired pressure, the device for compressing the gas having an inlet and an outlet, a heat exchanger device (1408) having a first inlet and a first outlet and a second inlet and a second outlet, a conduit connecting the first inlet and outlet and a conduit connecting the second inlet and outlet, said first and second conduits being in thermal contact along a major portion of their respective lengths, means (1406) for conveying the gas mixture from the outlet of the gas compression means to the first inlet of the heat exchanger means, at least one membrane device (1412) having an inlet for introducing a gas mixture to be separated, a gas separation membrane which divides the membrane device into a permeate chamber and a non-permeate chamber, said membrane being as specified in claim 1, a permeate gas outlet and a non-permeate gas outlet, means (1410) for conveying gas from the first outlet of the heat exchanger means to the inlet of the membrane means, a turbine expander (1424) for expanding a part or all of the permeate gas or non-permeate gas exiting the membrane means while cooling the gas, the turbine expander having an inlet and an outlet and being connected to at least partially drive a second means for compressing gas, a second device (1416) for compressing gas, having an inlet and an outlet, means (1422) for conveying a portion or all of the permeate gas or non-permeate gas from the membrane means to the inlet of the turbine expander, means (1414) for conveying a portion or all of the permeate gas or non-permeate gas from the membrane means to the inlet of the second means (1416) for compressing gas, and means for conveying gas from the outlet of the turbine expander (1424) to the second inlet of the heat exchanger means (1408). 56. Apparatus for carrying out the method of claim 1, comprising: a first device (1504, 1508) for compressing the gas mixture to a desired pressure, the first device for compressing the gas having an inlet and an outlet and optionally more than one stage with intermediate inlets, a heat exchanger device (1512) having a first inlet and first outlet, a second inlet and second outlet, and a third inlet and third outlet, a line connecting the first inlet and outlet, a line connecting the second inlet and outlet, and a line connecting the third inlet and outlet, said first, second and third lines are in thermal contact along a major portion of their respective lengths, a device (1510) for conveying the gas mixture from the outlet of the first device for compressing gas to the first inlet of the heat exchanger device (1512), a first membrane device (1516) having an inlet for introducing a gas mixture to be separated, at least one gas separation membrane which divides the membrane device into a permeate chamber and at least one non-permeate chamber, wherein this gas separation membrane is as specified in claim 1, a permeate gas outlet and a non-permeate gas outlet, a device (1514) for conveying gas which connects the first outlet of the heat exchanger device (1512) to the inlet of the first membrane device (1516), a second membrane device (1520) having an inlet for introducing gas, at least one membrane which divides the membrane device into at least one permeate chamber and at least one non-permeate chamber, a permeate gas outlet and a non-permeate gas outlet, means (1518) for conveying gas from the first membrane means (1516) to the second membrane means (1520), a turbine expander (1534) for expanding a portion or all of the gas exiting the non-permeate gas outlet of the first membrane means (1516) while cooling the gas, the turbine expander having an inlet and an outlet and being connected to at least partially drive a second means for compressing gas, a second device (1524) for compressing gas, having an inlet and an outlet, means (1522) for conveying gas from the non-permeate outlet of the first membrane means (1516) to the inlet of the turbine expander (1534), means (1536) for conveying gas from the outlet of the turbine expander (1534) to the second inlet of the heat exchanger means (1512), means (1540) for conveying gas from the second outlet of the heat exchanger means (1512) to either the inlet of the first means (1504, 1508) for compressing gas or an intermediate inlet for such a means, and means (1530) for conveying gas from the non-permeate outlet of the second membrane means (1520) to the third inlet of the heat exchanger means (1512).